HANDBOOK INDUSTRIAL

OF

MEMBRANE

TECHNOLOGY

Edited by

Mark C. Porter
Consultant Pleasanton, California

Reprint

Edition

El
"P

NOYES PUBLICATIONS
Westwood, New Jersey, U.S.A.

Copyright 01990 by Noyes Publications No part of this book may be reproduced any form or by any means, electronic or including photocopying, recording or by tion storage and retrieval system, without in writing from the Publisher. Library of Congress Catalog Card Number: ISBN O-8155-1205-8 Printed in the United States

or utilized in mechanical, any informapermission 88-17878

Published in the United States of America by Noyes Publications Fairview Avenue, Westwood, New Jersey 07675 10987654

NOTICE
To the best of the Publisher’s knowledge the information contained in this book is accurate; however, the Publisher assumes no responsibility nor liability for errors or any consequences arising from the use of the information contained herein. Mention of trade names or commercial products does not constitute endorsement or recommendation for use by the Publisher. Final determination of the suitability of any information, procedure, or product for use contemplated by any user, and the manner of that use, is the sole responsibility of the user. The book is intended for informational purposes only. Expert advice should be obtained at all times when implementation is being considered. xii

This roughly relates to the pore-size of the membrane. 0. and gas separations (GS). ultrafiltration (UF). With the exception of pervaporation. In addition. Even the most open MF membrane is capable of retaining yeast (3 to 12 microns) and tighter MF membranes can retain the smallest bacteria (Pseudomonas diminuta. UF and RO membranes. is about forty (40) microns in diameter. Little will be said concerning the use of membranes in medical applications as in artificial kidneys or for controlled drug release.l shows the pore sizes of MF. all particles or molecules larger than the rated pore size will be retained. which is just beginning to emerge as an industrial process. The notable exception to this is electrodialysis (ED) by which ions are separated under the influence of an electric field. Pressure driven processes include microfiltration (MF).08 microns. Most of the membrane processes are pressure driven.2). the chapter on coupled transport covers processes which are driven under the influence of a concentration gradient. The pressure driven liquid filtration processes (MF. The smallest particle which can be seen with the naked eye. Figure P. The most open UF membranes (almost always anisotropic) have a molecular weight cut-off of approximately one million (106) daltons which corresponds to about 0. There is some overlap in pore size because there are MF membranes (usually isotropic) available with pore sizes down to 0. under the best of lighting conditions.2 microns) (see Figure P. reverse osmosis (RO). A typical human hair has a diameter of eighty (80) microns. UF.1 microns (1000 8) is somewhat arbitrary.
V
. This means that membranes are filtering particles out of solution which are invisible to the naked eye. these MF membranes are used only for analytical applications and have no commercial importance for large scale processing. and RO) may be distinguished by the size of the particle or molecule the membrane is capable of retaining or passing. However.02 microns. Obviously. pervaporation (PV).Preface
This handbook emphasizes the use of synthetic membranes for separations involving industrial or municipal process streams. The dividing line between MF and UF membranes at 0. each of the above will be covered in a separate chapter.

we must exceed the osmotic several salts as a function NaCl will have an of

pressure as in Figure P.4. Hence, the name “reverse-osmosis.” osmotic water pressure concentration. close to 400 containing osmotic

psi (25 bar).

In practice,

pressures of 1,000

psi are used

in the reverse osmosis of sea water.

-T Osmotic

Head (pressure)

e

Pure Water
Semipermeable Membrane

Osmotic Equilibrium

Figure P.3:

Osmosis.

viii

Preface

Pressure .-

Semipermeable

Membrane

Reverse Osmosis Figure P.4: Reverse osmosis.

10,ooc

N

.c
1;

1,OOC

lO( I

IC)

0.1

I

1.0
Weight % solute

10.0

100.0

Figure P.5: Variation

of osmotic pressure with salt concentration.

Preface

ix

It has been market tions that treatment this potential is 1.6 in the

estimated

that

the

current field

worldwide along with

membrane/equipment New applicaexpanded to use in waste It is hoped the

billion emerging will

dollars

annually

(including

hemodialysis).

biotechnology make

promise volume

rapid growth

in the closing years of this century. contribution technology.

a significant

understanding

and limitations California

of membrane

Pleasanton, October

M.C. Porter Editor

1989

1 Synthetic Membranes and Their Preparation

Heiner Strathmann

INTRODUCTION In recent years, membranes and membrane separation techniques have grown from a simple laboratory tool to an industrial process with considerable technical and commercial impact.’ Today, membranes are used on a large scale to produce potable water from the sea by reverse osmosis, to clean industrial effluents and recover valuable constituents by electrodialysis, to fractionate macromolecular solutions in the food and drug industry by ultrafiltration,’ to remove urea and other toxins from the blood stream by dialysis in an artificial kidney, and to release drugs such as scopolamin, nitroglycerin, etc. at a predetermined rate in medical treatment.j Although membrane processes may be very different in their mode of operation, in the structures used as separating barriers, and in the driving forces used for the transport of the different chemical components, they have several features in common which make them attractive as a separation tool. In many cases, membrane processes are faster, more efficient and more economical than conventional separation techniques. With membranes, the separation is usually performed at ambient temperature, thus allowing temperature-sensitive solutions to be treated without the constituents being damaged or chemically altered. This is important in the food and drug industry and in biotechnology where temperature-sensitive products have to be processedB4 Membranes can also be “tailor-made” so that their properties can be adjusted to a specific separation task. Membrane science and technology is interdisciplinary, involving polymer chemists to develop new membrane.structures; physical chemists and mathematicians to describe the transport properties of different membranes using mathematical models to predict their separation characteristics; and chemical engineers to design separation processes for large scale industrial utilization. The most important element in a membrane process, however, is the membrane itself. To 1

2

Handbook of Industrial Membrane Technology

gain an understanding of the significance of the various structures used in different separation processes a brief discussion of the basic properties and functions of membranes, and the driving forces and fluxes involved is essential. Definition of a Membrane A precise and complete definition of a membrane which covers all its aspects is rather difficult, even when the discussion is limited to synthetic structures as in this outline. In the most general sense, a synthetic membrane is a barrier which separates two phases and restricts the transport of various chemical species in a rather specific manner.’ A membrane can be homogeneous or heterogeneous, symmetric or asymmetric in structure; it may be solid or liquid; it may be neutral, may carry positive or negative charges, or may be bipolar. Its thickness may vary between less than 100 nm to more than a centimeter. The electrical resistance may vary from several megohms to a fraction of an ohm, Mass transport through a membrane may be caused by convection or by diffusion of individual molecules, induced by an electric field, or a concentration, pressure or temperature gradient. The term “membrane”, therefore, includes a great variety of materials and structures, and a membrane can often be better described in terms of what it does rather than what it is. Some materials, though not meant to be membranes, show typical membrane properties, and in fact are membranes, e.g., protective coatings, or packaging materials. All materials functioning as membranes have one characteristic property in common: they restrict the passage of various chemical species in a very specific manner. Fluxes and Driving Forces in Membrane Separation Processes Separations in membrane processes are the result of differences in the transport rates of chemical species through the membrane interphase. The transport rate is determined by the driving force or forces acting on the individual components and their mobility and concentration within the interphase. The mobility is primarily determined by the solute’s molecular size and the physical structure of the interphase material, while the concentration of the solute in the interphase is determined by chemical compatibility of the solute and the interphase material, the solute’s size, and the membrane structure. The mobility and concentration of the solute within the interphase determine how large a flux is produced by a given driving force. In membrane separation processes, there are three basic forms of mass transport. The simplest form is the so-called “passive transport”. Here, the membrane acts as a physical barrier through which all components are transported under the driving force of a gradient in their electrochemical potential. Gradients in the electrochemical potential of a component in the membrane interphase may be caused by differences in the hydrostatic pressure, the concentration, the temperature, or the electrical potential between the two phases separated by the membrane. The second form of mass transport through the membrane interphase is the so-called “facilitated” transport. Here, the driving force for the transport of the various components is again the gradient in their electrochemical potential across the membrane. The different components, however, are coupled to a specific carrier in the membrane phase. Facilitated transport,

Synthetic Membranes and Their Preparation

3

therefore, is a special form of the passive transport. Completely different, however, is the third form of mass transport through membranes. It is generally referred to as “active” transport. Here, various components may be transported against the gradient of their electrochemical potential. The driving force for the transport is provided by a chemical reaction within the membrane phase. Active transport is coupled with a carrier in the membrane interphase and is found mainly in the membranes of living cells.6 It has, to date, no significance in synthetic membranes. The transport of mass in a membrane is a nonequilibrium process and is conventionally described by phenomenological equations such as Fick’s law which relates the fluxes of matter to the corresponding driving forces, i.e., a concentration gradient. The constant of proportionality is the diffusion coefficient. Driving forces in some membrane processes may be interdependent, giving rise to new effects. Thus, a concentration gradient across a membrane may result in not only a flux of matter, but, under certain conditions, also in the buildup of a hydrostatic pressure difference; this phenomenon is called osmosis. Similarly, a gradient in hydrostatic pressure may lead to a concentration gradient as well as a volume flow through the membrane; this phenomenon is called reverse osmosis. Frequently, fluxes of individual components are coupled, i.e., the flow of one component causes a flow of another.’ One example of the coupling of fluxes is the transport of bound water with an ion which is driven across a membrane by an electrical potential gradient. For membrane separation processes, only driving forces which induce a significant flux of matter are of practical importance. These driving forces are hydrostatic pressure, concentration, and electrical potential differences. These driving forces can also lead to the separation of chemical species. (a) A hydrostatic pressure difference between two phases separated by a membrane leads to a separation of chemical species when the hydrodynamic permeability of the membrane is different for different components. (b) A concentration difference between two phases separated by a membrane leads to a separation of various chemical species when the diffusivity and the concentration of the various chemical species in the membrane are different for different components. (c) A difference in the electrical potential between two phases separated by a membrane leads to a separation of various chemical species when the differently charged particles show different mobilities and concentrations in the membrane. The permeabilities of different components in a membrane depend on the mechanism by which the components are transported. For example, in homogeneous polymer membranes, the various chemical species are transported under a concentration or pressure gradient by diffusion. The permeability of these membranes is determined by the diffusivities and concentrations of the various components in the membrane matrix and the transport rates are, in general, relatively slow. In porous membrane structures, however, mass is transported under the driving force of a hydrostatic pressure difference via viscous flow and, in gen-

4

Handbook of Industrial Membrane Technology

eral, permeabilities are significantly higher than in diffusion-controlled membrane transport. In electrically charged membranes, usually referred to as ion-exchange membranes, ions carrying the same charge as the membrane material are more or less excluded from the membrane phase and, therefore, unable to penetrate the membrane.s The type of membrane and driving force required for a certain mass separation will depend on the specific properties of the chemical species in the mixture. For a given driving force, the flux through a unit membrane area is always inversely proportional to the thickness of the selective barrier. For economic reasons, membranes should in general be as thin as possible.

DISCUSSION OF TECHNICAL RELEVANT METHODS OF THEIR PREPARATION

SYNTHETIC

MEMBRANES

AND

Although synthetic membranes show a large variety in their physical structure and chemical nature, they can conveniently be classified in five basicgroups: (1) microporous media, (2) homogeneous solid films, (3) asymmetric structures, (4) electrically charged barriers and (5) liquid films with selective carriers. This classification, however, is rather arbitrary and there are many structures which would fit more than one of the abovementioned classes, e.g., a membrane may be microporous, asymmetric in structure, and carry electrical charges. Any other classification of synthetic membranes, e.g., according to their application or methods of preparation, would serve the same purpose of phenomenologically categorizing the various types of synthetic membranes. Neutral Microporous Membranes The neutral, microporous films represent a very simple form of a membrane which closely resembles the conventional fiber filter as far as the mode of separation and the mass transport are concerned. These membranes consist of a solid matrix with defined holes or pores which have diameters ranging from less than 2 nm to more than 20 pm. Separation of the various chemical components is achieved strictly by a sieving mechanism with the pore diameters and the particle sizes being the determining parameters. Microporous membranes can be made from various materials, such as ceramics, graphite, metal or metal oxides, and various polymers. Their structure may be symmetric, i.e., the pore diameters do not vary over the membrane cross section, or they can be asymmetrically structured, i.e., the pore diameters increase from one side of the membrane to the other by a factor of 10 to 1,000. The properties and areas of application of various microporous filters are summarized in Table 1 .I. Symmetric Microporous Sintered Membranes. Sintered membranes are the simplest in their function and in the way they are prepared. The structure of a typical sintered membrane is shown in the scanning electron micrograph of Figure 1.1. This photograph shows a microporous membrane made out of polytetrafluoroethylene by pressing a fine powder into, a film or plate of 100 to 500 pm thickness and then sintering the film at a temperature which is just below the melting point of the polymer. This process yields a microporous structure of relatively low porosity, in the range of IO to 40%. and a rather irregular pore structure with a very wide pore size distribution.

Synthetic

Membranes

and Their

Preparation

5

Table

1.1:

Microporous

Membranes,

Their

Preparation

and Application
Application

Membrane

type

Membrane

material

Pore size

Manufacturing

process

Ceramic, metal or polymer powder
Mlcroporous membrane

0.1 -20

pm

pressIng

and 51nterlng

Mlcrollltration

of powder

Homogeneous

polymer

0.' -10

I'm

stretching polymer

of sheet.

extruded

Microfiltration, burn dressings, artificial blood vessels

sheets (PE, PTFE)

Homogeneous sheets (PC)

polymer

0.02 -10

I'm

track-etching

Microfiltration

Polymer (CN. CA)

solution

0.01 -,

"111

phase inversion

Microfiltration. ultrafiltration. sterilization

2nI11-'1'I11

Figure powder.

1.1:

SEM

of a microporous

sintered

membrane

prepared

from

a PTF E-

Sintered membranes are made on a fairly large scale from ceramic materials, glass, graphite and metal powders such as stainless steel and tungsten.' The particle size of the powder is the main parameter determining the pore sizes of the final membrane, which can be made in the form of discs, candles, or fine-bore tubes. Sintered membranes are used for the filtration of colloidal solutions and suspensions. This type of membrane is also marginally suitable for gas separation. It is widely used today for the separation of radioactive isotopes, especially uranium.

6

Handbook

of

Industrial

Membrane

Technology

Stretched Membranes. Another relatively simple procedure for preparing microporous membranes is the stretching of a homogeneous polymer film of partial crystallinity. This technique is mainly employed with films of polyethylene or polytetrafluoroethylene which have been extruded from a polymer powder and then stretched perpendicular to the direction of extrusion.l0,11 This leads to a partial fracture of the film and relatively uniform pores with diameters of 1 to 20 .urn. A typical stretched membrane prepared from tetrafluoroethylene is shown in the scanning electron micrograph of Figure 1.2.

Figure

1.2:

SEM of a microporous perpendicular

membrane

prepared

by stretching

an extruded

PTFE.film

to the direction

of extrusion.

These membranes, which have a very high porosity, up to 90%, and a fairly regular pore size are now widely used for microfiltration of acid and caustic solutions, organic solvents, and hot gases. They have to a large extent replaced the sintered materials used earlier in this application. Stretched membranes can be produced as flat sheets as well as tubes and capillaries. The stretched membrane made out of polytetrafluoroethylene is frequently used as a water repellent textile for clothing, such as parkas, tents, sleeping bags, etc. This membrane type has, because of its very high porosity, a high permeability for gases and vapors, but, because of the hydrophobic nature of the basic polymer, is up to a certain hydrostatic pressure completely impermeable to aqueous solutions. Thus, the membrane is repellent to rain water but permits the water vapor from the body to permeate. More recently, this membrane has also been used for a novel process, generally referred to as membrane distillation, i.e., to remove ethanol from

Synthetic Membranes and Their Preparation

7

fermentation broths or wine and beer to produce low alcohol products” and for desalination of seawater. These membranes are also used for desalination of saline solutions and in medical applications such as burn dressings and artificial blood vessels. Capillary Pore Membranes. Microporous membranes with very uniform, nearly perfectly round cylindrical pores are obtained by a process generally referred to as track-etching.13 The membranes are made in a two step process. During the first step, a homogeneous 10 to 15 pm thick polymer film is exposed to collimated, charged particles in a nuclear reactor. As particles pass through the film, they leave sensitized tracks where the chemical bonds in the polymer backbone we broken. In the second step, the irradiated film is placed in an etching bath. In this bath, the damaged material along the tracks is profcrentially etched forming uniform cylindrical pores. The entire process is schematically shown in Figure 1.3. The pore density of a track-etched membrane is determined by the residence time in the irradiator, while the pore diameter is controlled by the residence time in the etching bath. The minimum port diameter of theso mcmbrancs is approximately 0.01 pm. The maximum pore size that can bc achieved in tracketched membranes is determined by the etching procedure. The polymer will not only be dissolved along the sensitized track left by tho penetrating particle but also on both surfaces of the film. Thus, with exposure time in the etching medium the pore sizes increase and the thickness of the film is correspondingly reduced. The scanning electron micrograph in Figure 1.4 shows a typical tracketched polycarbonate membrane. Capillary pore membranes are prepared today mainly from polycarbonato and polyester films. The advantage of these polymers is that they are commercially availabla in very uniform films of 10 to 15 pm thickness which is the maximum penetration depth of collimated particles obtained from a nuclear reactor which have an oncrgy about 0.8 to 1 MeV. Particles with higher enorgy, up to IO McV, may be obtained in an accelerator. They are used today to irradiate thickor polymer films, up to 50flm thickness, or inorganic materials such as mica.14 Wowever, these mumbrancs arc not yet available on a commercial basis.

Because of their narrow pore size distribution and low tendency to plug, capillary pore membranes made from polycarbonate and polyester have found application on a large scale in analytical chemistry and microbiological laboratories, and in medical diagnostic procedures.15 On an industrial scale, capillary pore membranes are used for the production of ultrapure water for the electronic industry. Here, they show certain advantages over other membrane products because of their short "rinse down" time and good long-term flux stability. Because of their surface filter characteristics, particles retained by the membranes can be further monitored by optical or scanning electron microscopy. Figure 1.5 shows a scanning electron micrograph of asbestos fibers accumulated on a capillary pore membrane in an air pollution control application. The membranes are also used in standard clinical tests for red blood cell deformability studies.

Figure 1.5: SEM pore membrane .

of asbestos

filter

accumulated

on the

surface

of a capillary

Synthetic

Membranes

and Their

Preparation

9

Human red blood cells have a diameter of approximately 6 to 8 Jlm. The human body, however, contains capillaries approximately 3 Jlm in diameter. To pass through these vessels the blood cells have to deform correspondingly. Healthy cells will do this readily but malignant cells will not. By filtering blood through a 3 Jlm capillary pore membrane certain blood deficiencies can be monitored.16 Symmetric Microporous Phase Inversion Membranes. The most important commercially available, symmetric, microporous membranes are prepared by the so-called phase inversion process.17 In this process, a polymer is dissolved in an appropriate solvent and spread as a 20 to 200 Jlm thick film. A precipitant such as water is added to this liquid film from the vapor phase, causing separation of the homogeneous polymer solution into a solid polymer and a liquid solvent phase. The precipitated polymer forms a porous structure containing a network of more or less uniform pores. A microporous cellulosic membrane made by phase inversion is shown in the scanning electron micrograph of Figure 1.6. This type of membrane can be made from almost any polymer which is soluble in an appropriate solvent and can be precipitated in a nonsolvent.18 By varying the polymer, the polymer concentration, the precipitation medium, and the precipitation temperature, microporous phase inversion membranes can be made with a very large variety of pore sizes (from less than 0.1 to more than 20 Jlm) with varying chemical, thermal, and mechanical properties. These membranes were originally prepared from cellulosic polymers by precipitation at room temperature in an atmosphere of approximately 100% relative humidity.19 Lately, symmetric microporous membranes are also prepared from Nylon 66, Nomex, polysulfone, and polyvinylidene difluoride by precipitation of a cast polymer solution in aqueous liquid.20

Figure prepared

1.6:

SEM

of the

surface from

of a microporous a homogeneous

cellulose

nitrate

membrane vapor

by precipitation

polymer

solution

by water

precipitation.

mal gelation?l It can be applied to metal alloys and glasses as well as polymer solutions. solvent but merely by cooling the solution to a point where a two phase system forms. e.7:
SEM by
of
the
surface gelation
of
a microporous a hot
polypropylene polymer
membrane solution. The precipitation of the polymer. particularly well suited when a complete re-
.. however. and the detoxification of blood in an artificial kidney. microporous polymer membranes made by phase inversion are widely used for separations on a laboratory and industrial scale?2 Typical applications range from the clarification of turbid solutions to the removal of bacteria or enzymes.
precipitated
thermal
from
homogeneous
The symmetric. The resulting open foam structure is shown in Figure 1. therefore. solvent system. since these polymers are not readily dissolved at room temperature. I n addition to the simple "sieving" effect. This membrane preparation technique is usually referred to as ther. A solution of 20 to 30% polymer is spread at elevated temperature into a film. solution temperature. microporous phase inversion membranes often show a high tendency of adsorption because of their extremely large internal surface. is dissolved in an appropriate amine at elevated temperatures. The separation mechanism is that of a typical depth filter which traps the particles somewhere within the structure. the detection of pathological components.'
Figure
1.7. The pore size depends on polymer concentration. is not induced by the addition of a non. and cooling rate. They are.g. Polypropylene.10
Handbook
of Industrial
Membrane
Technology
Polypropylene or polyethylene can also be used for the preparation of microporous membranes. However. the preparation technique must be slightly varied.

which diagram
24 to 48 hours. structure simple: two difwith types of glass are homogeneously base leaching. procedure mixed.8: SEM
schematic diagram
of cultured
microorganisms
in wotcr
quality
control
tests and
of the process.% pore and can defined tubes. immobilization technology. of their metal acid or alloys.
placed on a nutrient the microorganisms assist in counting the
pad in an incubator colony density. of Figure
The process is shown in
and the photograph
Figure 1. quality which control retains They
Membranes
and Their
Preparation
11
such as viruses or bacteria
is desired.23 for
or even microorganisms used for culturing the water The Grid membrane During marked
to be used in modern of microorganisms through is then
are also widely
In this test. Microporous have found their
membranes leaching
can be prepared
alloys such as Ni-Al-Cr
by subsequent main application
of one component. one glass type is removed is obtained to several nanometers.
grow into easily visible colonies. The preparation Thus.8. such as flat because from These importance
a microporous in various are metal of
sizes in the range of a few angstroms be made which hollow fibers.
is an indication
of water quality.
membranes
in gas separation
processes. is relatively then.
is filtered this time. While
. particular
Porous
glass membranes
configurations.Synthetic moval of components.
high surface area. the schematic
membranes 1.
Symmetric Microporous Phase Inversion Membranes from Inorganic Materials.
a membrane
all microorganisms. Microporous membranes can also be prepared by phase inversion from
glass or metal ferent by well sheet.
They
are suited for bioin water
of enzymes tests.

e. the two basic properties required of any membrane.12
Handbook of Industrial Membrane Technology
the microporous metal membranes are mainly used for gas separation tasks. The porous sublayer serves only as a support for the thin and fragile skin and has little effect on separation characteristics or the mass transfer rate of the membrane. since here the unique properties in terms of high mass transfer rates and good mechanical stability can be best utilized.9.9: Schematic drawing of the cross section of a (a) symmetric and (b) asymmetric membrane. i. high mass transport rates for certain components and good mechanical strength. are physically separated. Conventional symmetric structures act as depth filters and retain particles within their internal structure.. asymmetric membranes are most fouling resistant. Two techniques are used to prepare asymmetric membranes: one utilizes the phase inversion process and the other leads to a composite structure by depositing an extremely thin polymer film on a microporous substructure. or gas separation.26
a)
b)
Figure 1. Asymmetric membranes are surface filters and retain all rejected materials at the surface where they can be removed by shear forces applied by the feed solution moving parallel to the membrane surface.1 to 1 pm) selective skin layer on a highly porous (100 to 200 pm) thick substructure. The difference in the filtration behavior between a symmetric and an asymmetric membrane is shown schematically in Figure 1. Its separation characteristics are determined by the nature of the polymer and the pore size while the mass transport rate is determined by the membrane thickness.10. such as reverse osmosis. The very thin skin represents the actual membrane. In this membrane.
. Asymmetric membranes are used primarily in pressure driven membrane processes. and in artificial kidneys for detoxification of blood streams. ultrafiltration. as indicated in the schematic drawing of Figure 1. ultrafiltration of wastewater. since the mass transport rate is inversely proportional to the thickness of the actual barrier layer.25
Asymmetric Microporous Membranes The most important membrane used today in separation processes is composed of a rather sophisticated asymmetric structure.
In addition to high filtration rates. which shows the cross section of (a) a symmetric and (b) an asymmetric membrane. microporous glass membranes have been used mainly for separation of liquid mixtures including desalination of sea and brackish waters.. These trapped particles plug the membrane and the flux declines during use. An asymmetric membrane consists of a very thin (0.

which provided the necessary structural information. Detailed descriptions of membrane preparation techniques are given in the literature.*s The Formation Mechanism of Microporous Symmetric or Asymmetric Membranes.2J~30 Only after extensive use of the scanning electron microscope. Generally. the asymmetric structure was obtained mainly in membranes made from cellulose acetate.
(2) The solution is cast into a film of typically 100 to 500pm thickness.. Asymmetric phase inversion membranes can be prepared from cellulose acetate and many other polymers by the following general preparation procedure:*’
(I) A polymer is dissolved in an appropriate solvent to form a solution
containing 10 to 30 weight % polymer. These membranes were made from cellulose acetate and yielded fluxes 10 to 100 times higher than symmetric structures with comparable separation characteristics. and a solvent-rich liquid phase. where precipitation occurs first and most rapidly. e. which leads to the asymmetric membrane structure. which forms the liquid-filled membrane pores.
--_
-
characteristics of a (a) sym-
Preparation Procedures of Asymmetric Membranes. was it possible to rationalize the various parameters for membrane preparation processes. The development of the first asymmetric phase inversion membranes was a major breakthrough in the development of ultrafiltration and reverse osmosis. which forms the membrane structure.
(3) The film is quenched in a nonsolvent typically water or an aqueous solution. But later it became ap-
. At first. are much smaller than those in the interior or the bottom side of the film.10: Schematic diagram of the filtration metric and (b) asymmetric membrane. There are different variations to this general preparation procedure described in the literature.Synthetic Membranes and Their Preparation
13
a)
bP
-Figure 1. the pores at the film surface. During the quenching process the homogeneous polymer solution separates into two phases: a polymer-rich solid phase. Loeb and Sourirajan used an evaporation step to increase the polymer concentration in the surface of the cast polymer solution and an annealing step during which the precipitated polymer film is exposed for a certain time period to hot water of 70” to 80°C. The original recipes and subsequent modifications for preparing asymmetric membranes are deeply rooted in empiricism.g.

. e. Phase inversion membranes can be prepared from any polymer mixture which forms. such as the chemical potentials and diffusivities of the individual components and the Gibb’s free energy of mixing of the entire system. however. During a phase inversion process. because of their poor solubility.. which Kesting*’ refers to as phase inversion. precipitation of a simple two-component polymer-solvent casting solution can be brought about by imbibing precipitant from the vapor phase. If the liquid phase is discontinuous. But in all cases. under certain conditions of temperature and composition.N-bis-(2-hydroxyethyl)tallow amine. causing precipitation as the system becomes enriched in precipitant. metals or glassesseem superficially very different. a polymer powder will be obtained instead of a rigid structure. a homogeneous solution and separates at a different temperature or composition into two phases. The Phase Separation Process and Its Relation to the Formation Mechanism of Microporous Membranes Detailed recipes given in the literature for the preparation of microporous structures from polymers. (2) of a homogeneous solution of two or more com-
Evaporation of a volatile solvent from a homogeneous solution of two or more components. The actual phase separation can be induced by several means. Yet. polypropylene dissolved in N. One technique described in the literature is based on the controlled evaporation of a volatile solvent from a three-component mixture of solvent/precipitant/polymer. as indicated earlier. This technique was used by Zsigmondy3’ and more recently by Kesting3* for the preparation of ultrafiltration and reverse osmosis membranes. This process is called thermal gelation. another technique is to bring about precipitation by cooling a casting solution which forms a homogeneous solution only at elevated temperature. would otherwise be inaccessible to the phase inversion membrane preparation techniques. This technique was the basis of the original microporous membranes and is still used commercially today by several companies. A further condition. a onephase polymer solution is converted into a two-phase system consisting of a solid (polymer-rich) phase which forms the membrane structure and a liquid (polymer-poor) phase which forms the pores in the final membrane. the basic membrane formation mechanism is governed by similar thermodynamic and kinetic parameters. Thermal gelation is not only applicable to polymers which.33 Identification and description of the phase separation process is the key to understanding the membrane formation mechanism-a necessity for optimizing membrane properties and structures. is that both phases are continuous. The techniques used to induce the phase separation in a homogeneous solution for the preparation of microporous membranes can be related to three basic procedures: (1) Thermogelation ponents. it can also be used for making microporous membranes from glass mixtures and metal alloys in combination with a leaching procedure. Alternatively. a closed-cell foam structure will be obtained and if the solid phase is discontinuous.g.14
Handbook of Industrial Membrane Technology
parent that the process of making skin-type asymmetric membranes by precipitating a polymer solution is just a special case of a general procedure.

11. which shows a phase diagram of a two component mixture of a polymer and a solvent as a function of temperature. which at sufficiently high temperature forms a homogeneous solution for all compositions. Phase separation due to thermal gelation.
15
All three procedures may result in symmetric microporous structures or in asymmetric structures with a more or less dense skin at one or both surfaces suitable for reverse osmosis. polymer and solvent respectively. This behavior is illustrated schematically in Figure 1. but at a lower temperature shows a miscibility gap over a wide range of compositions. The simplest procedure for obtaining a microporous system is by thermogelation of a two component mixture. and points on the line P-S describe mixtures of these two components.Synthetic Membranes and Their Preparation (3) Addition of a nonsolvent or nonsolvent mixture to a homogeneous solution. at a temperatureT.34 Phenomenological Description of Phase Separation. the system must have a miscibility gap over a defined concentration and temperature range.
. If a homogeneous mixture of the composition X.ll: Schematic diagram showing the formation of a microporous system by thermal gelation of a two-component mixture exhibiting a miscibility gap at certain conditions of temperature and composition.. Solution
/
Miscibility
Gap
\
Solid Phase \
-
Composition
Y
\
Liquid Phasi
Figure 1 . evaporation of solvent and addition of nonsolvent can be illustrated with the aid of the phase diagram of a polymer solution. Homogeneous . ultrafiltration or microfiltration. that is. Thermogelation of a Two-Component Polymer Mixture. The only thermodynamic presumption for all three preparation procedures is that the free energy of mixing of the polymer system under certain conditions of temperature and composition is negative.
The points P and S represent the pure components.

. the system is separated into two phases: a polymer-rich phase. The corners of the triangle represent the pure components. Polymer
Solvent
Non-solvent
Figure 1. The lines B’-B and B”-B represent the ratio of the amounts of the two phases in the mixture. the composition of the mixture will change from that represented by point A to that represented by point B. Boundary lines between any two corners represent mixtures of two components.12: Schematic diagram showing the formation of a microporous system by evaporation of a solvent from a three-component mixture exhibiting a miscibility gap at certain conditions of temperature and composition.16
Handbook of Industrial Membrane Technology
as indicated by the point A in Figure 1 . polymer-poor liquid phase. whereas in a different range-the miscibility gap-the system separates into two distinct phases.ll. all three components are completely miscible. This process utilizes a three-component mixture: a polymer. as indicated in Figure 1. At point B. If the volatile solvent is completely evaporated from a homogeneous mixture of 10% polymer. 60% solvent and 30% nonsolvent. and the phase B” forming the liquid filled pores of the membrane. Since this point is situated within the miscibility gap. indicated by point B’ forming the rigid structure. is cooled to the temperature Tz. as indicated by point A in Figure 1. Evaporation of a Volatile Solvent from a Three-Component Polymer Solution. which by itself is a nonsolvent for the polymer. as indicated by point 6. the overall porosity of the obtained microporous system. This three-component mixture is completely miscible over a certain composition range but exhibits a miscibility gap over another composition range. it will separate into two different phases.12. which represents an isothermal phase diagram of the three components. the system consists of only two components: polymer and nonsolvent. Within a certain compositionally defined range of thermodynamic states.12. the composition of which are indicated by the points B’ and 6”. a volatile solvent and a third component. and any point within the triangle represents a mixture of all three components. that is. The point 6’ represents the polymer-rich solid phase and the point 6” the solvent-rich.

This technique can again be rationalized with the aid of a three-component isothermal phase diagram shown schematically in Figure 1. It predicts under what conditions of tem-
. If a nonsolvent is added to a homogeneous solution consisting of polymer and solvent. Further exchange of solvent and nonsolvent will lead to the final composition of the membrane.13.
The description of the formation of microporous systems by means of the phase diagrams.13.13.Synthetic Membranes and Their Preparation
17
Addition of a Nonsolvent to a Homogeneous Polymer Solution. the polymer concentration in the polymer-rich phase will be high enough to be considered as solid. if the solvent is removed from the mixture at about the same rate as the nonsolvent enters. and.13: Schematic diagram showing the formation of a microporous system by addition of a nonsolvent to a homogeneous polymer solution in a threecomponent mixture exhibiting a miscibility gap at certain conditions of temperature and composition. the membrane structure is more or less determined. At this point. This phase diagram of the three-component mixture exhibits a miscibility gap over a wide range of compositions. the composition of the system will reach the miscibility gap and two separate phases will begin to form: a polymer-rich phase represented by the upper boundary of the miscibility gap and a polymer-poor phase represented by the lower boundary of the miscibility gap. Point B represents a mixture of the solid polymer-rich phase and the liquid solvent-rich phase as represented by points B’ and B” respectively. At a certain composition of the three-component mixtures. the composition of which is indicated by the point A on the solvent-polymer line. as illustrated in Figure 1.11 to 1.
Polymer
Solvent /
Non-solvent
Figure 1. This composition is represented by point D in Figure 1. the porosity of which is determined by point B. the composition of the mixture will change following the line A-B. is based on the assumption of thermodynamic equilibrium. At point C.

In this state. viscosity. i..18
Handbook of Industrial Membrane Technology
perature and composition a system will separate into two phases and the ratio of the two phases in the heterogeneous mixture. whether the membrane has a symmetric or asymmetric structure or a dense skin at the surface.e. Equilibrium thermodynamics is not able to offer any explanation about structural variations within the membrane cross section. are difficult to determine by independent experiments and. which are determined by the spatial distribution of the two phases. the diffusion coefficient. In this state. i. The transport processes are determined by thermodynamic parameters.T = const) (1)
(2) An unstable state where the homogeneous solution separates spontaneously into two phases which are then in equilibrium. which allows rationalization of the membrane formation and correlation of the various preparation parameters with membrane structures and properties. are not readily available. However. In this state. which is located within the miscibility gap. However. the overall porosity.. transport of individual components has to take place.. the viscosity of the solution and the chemical potential gradients which act as driving forces for diffusion of the various components in the mixture. which constitutes the actual membrane formation process. The chemical potential and diffusivities of the various components in the system. etc.e.40 At constant pressure and temperature. which depend on system properties. the free energy of mixing is zero AG = 0 (3) (2)
The free energy of mixing of a system describes the thermodynamic state of the system and thus provides information about the system stability.35”9 Mathematical Description of Phase Separation. three different states can be distinguished: (1) A stable state where all components are miscible in a single phase. This makes a quantitative description of the membrane formation mechanism nearly impossible. however. is possible. temperature. Because these parameters change continuously during the phase separation. Fick’s law relates the diffusion coefficient to concentration gradients. frozen states will often be obtained that are far from equilibrium and that can be stable for long time periods. such as the diffusivities of the various components in the mixture. These parameters are determined by kinetic effects. no information is provided about the pore sizes. Especially in poiymer systems. that is. The thermodynamic state of a system of two or more components with limited miscibility can be described in terms of the free energy of mixing. and their dependencies on composition. the free energy of mixing is negative AC<0 (3) An equilibrium state given by the phase boundary composition. If a system is unstable and separates in two coexisting phases. and by kinetic parameters. which are expressed by driving forces. no transient states of equilibrium will be achieved. the actual driving forces for any mass transport are gradients in the chemical po-
. which are determined by diffusivities. A qualitative description. the free energy of mixing is positive AC>0 (P. therefore.

(6G)
PT = (Zi). since the diffusion coefficient is directly related to the chemical potential of the individual components.
j . which is always positive.Synthetic Membranes and Their Preparation
19
tential of the individual components. (
(Ix5
6 IllXi
>
> 0
(5)
(2)
Unstable state < 0
and
oi =. 6In C
. The stable. y
.. i. The diffusion coefficients of individual components can be related to their chemical potential driving force by:41
0. Equations (4) to (7) indicate that in the case of spontaneous demixing the diffusion coefficient will become negative. which again is a function of the Gibb’s free energy.P
and
oi. R is the gas constant and T is the absolute temperature. negative or disappearing values of the free energy of mixing but. = Bi
I
( )
G
bPi
(4)
#
PT
Here.. . Since: Pi = Pp
l
PT lnXi
l
RTln
ty
(8)
6
RTInXi 6Xf
l
6 RT In r: dxi
(9)
=
xy RT
1
l
6 Inf:
6 InXi
(IO)
= Pi
.
y
I
. Bi is a mobility term. the individual components will diffuse against their concentration gradients. Di is the diffusion coefficient of component i.e. I_ciis its chemical potential and Xi its mol fraction.
(II)
.
. Bi is a mobility term and Xi the molfraction of the component i.
l
InXi
<
o
(6)
(3)
Equilibrium state
=
0
and
pi =
7
Bi RT I
63
bltlXi
=o
(7)
Here fis is an activity coefficient referring to the pure phase. unstable and equilibrium states of the system are not only characterized by positive. the state of a system is also characterized by: (1) Stable state
6P ( 70
ax 1 T.

The membrane shows a “sponge”like structure with no skin on the bottom or top surface. the four different structures shown in Figures 1. The membrane shows a sponge-type structure with a dense skin at the surface and a porous structure underneath with a relatively uniform pore size distribution over the entire cross section. Many aspects of the formation of symmetric or asymmetric membranes can be rationalized by applying the basic thermodynamic and kinetic relations of phase separation. Photograph (c) shows the cross section of a typical reverse osmosis membrane prepared from a polyamide solution and precipitated in water by immersing the polymer solution into a water bath. Using scanning electron microscope techniques.I 5 which shows scanning electron micrographs of the structures of two phase inversion membranes. This structure is often obtained from casting solutions with relatively high polymer content.15 (a) shows a sponge-type structure consisting of spherical cells. The membrane shows a “finger-“-type structure with a dense skin at the surface and large pores penetrating the entire membrane cross section. which in the literature is called “nodular”. Typical Membrane Structures. Photograph (a) shows a cross section of a microfiltration membrane prepared from a cellulose nitrate solution by precipitation in a humidity controlled environment.20
Handbook
of industrial
Membrane Technology
The chemical potentials of the individual components and their dependence on concentration of other species in a mixture and the mobility term I3 and its dependence on viscosity determine the phase separation process. Membranes with this type of structure can usually be dried without changing their mass transport properties. Photograph (d) shows the cross section of a reverse osmosis membrane prepared from a polyamide solution and precipitated in a water-solvent mixture. Figure 1 . A mathematical treatment of the formation of porous structures is difficult. There are. It consists of small polymer beads randomly fused together. sol and gel structures-which are not directly related to the thermodynamics of phase separation but which will have a strong effect on membrane structures and properties. The pores increase in diameter from the top to the bottom side. General Observations Concerning Structures and Properties of Phase Inversion Membranes. The membrane shows a sponge-type structure with a dense skin at the surface and a porous structure underneath with increasing pore diameter from the top to the bottom side of the membrane.14 (a) to (d) can be observed for phase inversion membranes. polymer relaxation. Looking at the fine structure of a membrane very often two types of structures may be obtained as indicated in Figure 1 . Photograph (b) shows a cross section of a typical ultrafiltration membrane prepared from a polyamide solution and precipitated in water by immersing the polymer solution into a water bath. other parameters-such as surface tension. however. Before going into any detailed discussion of the formation mechanism of microporous membranes. Figure 1. preparation procedures. The only differences in the preparation procedures of the membranes shown in photographs (b) and (c) are the polymer concentration and the precipitation temperature. The structure is often obtained from casting solutions with low poly-
. and mass transport properties are described. But many aspects of membrane structures and the effect of various preparation parameters can be qualitatively interpreted. several general observations concerning the membrane structure.15 (b) shows a structure.

(c) asymmetric membrane with a “sponge’‘-type structure.14: Scanning electron micrograph of membrane cross sections with typical structures: (a) symmetric microporous membrane without a “skin”.
b)
c)
d)
Figure 1. Very often in asymmetric membranes the structure may change from nodular to “cellular” type over the membrane cross section. as shown in Figure 1 .Synthetic Membranes and Their Preparation
21
mer concentrations. (b) asymmetric membrane with a “finger”-type structure and a dense skin at the surface.I 6. a dense skin and a uniform pore size distribution in the substructure.
. (d) symmetric membrane with a sponge structure. and pore sizes increasing from the surface to the bottom side. a dense skin.

4. The flux and retention properties. The overall porosity also decreases and the gelation time increases with increasing polymer concentration. These membranes show completely different structures and filtration properties which are listed in Table 1. The polymer concentration is a very significant parameter for tailoring a membrane in terms of its structure and separation properties.Synthetic Membranes and Their Preparation
23
Figure 1.3 indicate that the same type of membrane can be prepared from various polymers and that from one polymer-solvent system. The scanning electron micrographs of Figures 1.16: Scanning electron micrograph of a membrane showing a nodular structure at the surface which converts into a cellular structure towards the bottom side of the membrane. as shown in Table 1.2) prepared from three different polymers (cellulose acetate. The scanning electron micrograph of Figures 1. With increasing polymer concentration.
.18 (a) to (c) show the cross sections of membranes made from one polymer-solvent system (polyamide in DMAC) with different polymer concentrations. This is demonstrated in Figure 1. the structure changes from a typical finger structure (5% polyamide in the casting solution) to a typical sponge structure (22% polyamide in the casting solution). the permeability decreases and the retention increases. various types of membranes can be made. the membrane porosity and the rate of precipitation change in a corresponding pattern.I 7 and 1 .
The Effect of the Polymer and Polymer Concentration on Membrane Structures and Properties.2 and 1.19 where the scanning electron micrographs of membranes are shown which are prepared from casting solutions of different polyamide concentrations in NMP. With increasing polymer concentration. The scanning electron micrographs of Figures 1 .I8 and the data listed in Tables 1. polyamide and polysulfone) by precipitation in a water bath.17 (a) to (c) show the cross sections of three membranes with nearly identical structures and ultrafiltration properties (listed in Table 1.3.

26
Handbook of Industrial Membrane Technology
Figure 1.19: Scanning electron micrographs of membrane crosssections prepared from various polyamide concentrations in NMP by precipitation in water at room temperature.
.

the tendency for a change from a sponge to a finger
. MgSO4 chrome C bov. The rate of precipitation and its relation to the membrane structure can be easily observed with an optical microscope or by other means.42 The Effect of the Polymer-Solvent-Precipitant System on Membrane Structure and Properties. the pore diameters are inversely proportional to the rate of precipitation. and lower precipitation rates lead to coarser structures. Certain membrane structures can be correlated with the rate of precipitation.Synthetic Membranes and Their Preparation Table 1.4: Rate of Precipitation and Filtration Properties of Membranes Prepared From Polyamide-NMP Solutions of Various Concentrations by Precipitation in Water at Room Temperature
Polymer concentration (%)
5x 0 0 10
27
Rejection’ Cyto-. Higher precipitation rates lead to finer pore structures. The polymer-solvent-precipitant interaction can be approximately correlated in terms of the disparity of the solubility parameter of polymer and solvent. albumin Filtration rate km/r
56
Celation Porosity time++ (5) 32 40 52 83 142 212
x 10’)
(vol. Very high precipitation rates (short gelation times) always lead to a “finger” structure. (1) The smaller the solubility parameter disparity of solvent and polymer.6
85 81 79 77 76
+
++ x xx
the rejection was determined with solutions of I % solids
for
entire membrane
applied pressure 5 bar applied pressure LOObar
Rate of Precipitation and Membrane Structure. The solvent and the precipitant used in membrane preparation determine both the activity coefficient of the polymer in the solvent-precipitant mixture and the concentration of the polymer at the point of precipitation and solidification. the better the compatibility of solvent and polymer. and very slow precipitation rates very often lead to symmetric membranes with no defined skin at the surface. while slow precipitation rates (long gelation times) lead to asymmetric membranes with a “sponge” structure. The effect of the polymer on membrane structure and properties is closely related to the solvent used in the casting solution and the precipitant.%)
91
IO x IS I 18 xx 20 xx 22 xx
0 8 7s 90 98
43 92 100 100 LOO
84 100 100 100 100
32 9 18 4 1. a structure with a very uniform pore size distribution over the entire cross section of the membrane. Thus. If a sponge-type structure is obtained. and the slower the precipitation of the polymer.

(2) The higher this disparity.. The tendency to change from a sponge to a finger structure will thus increase with decreasing compatibility of polymer and precipitant. This is demonstrated in the scanning electron micrograph of Figure 1. the less compatible are polymer and precipitant. i .20: Scanning electron micrographs of membrane cross sections prepared from 15% polyamide in DMAc precipitated in water-glycerin mixture at room temperature. and the faster will be the precipitation.e. the disparity in the solubil ity parameter of polymer-water is larger than of polymer-glycerin. Since the compatibility of the polymer with glycerin is slightly better than that with the water. Additives to the casting solution or the precipitation bath have a similar effect.28
Handbook of Industrial
Membrane Technology
structure membrane increases with decreasing compatibility of solvent and polymer..rln
Figure 1.
. the tendency of the membrane structure to go from a finger to a sponge structure will increase with increasing glycerin content in the precipitation bath.36
80%Gly.20 which shows the cross section of membranes prepared from a solution of 15% polyamide in DMAc precipitated in water-glycerin mixtures of different compositions.

the polymer matrix breaks at weak spots. as indicated in the scanning electron micrograph of Figure 1. At the point of first phase separation. Because this process is so slow. bulk movement of the precipitated polymer takes place to form the matrix of the final membrane. the precipitant is introduced from the vapor phase. and a more or less homogeneous structure is obtained without a dense skin on the top or bottom side of the polymer film. In this case. polymer-rich phase undergoes continuous desolvation. Traveling from the glass plate towards the precipitation bath these layers are: (1) The casting solution layer: This layer is closest to the glass plate. (2) The fluid polymer layer: This layer lies between the point of precipitation and the point of solidification. Instead. In the first method. Because the polymer is solid. the casting solution can be divided into three layers. This structure can be understood when the concentration profiles of the polymer. the membrane precipitates rather slowly and in the same way over the entire film cross section. the polymer-rich phase still contains a high solvent concentration and a low precipitant concentration and is still liquid. The polymer nearest the precipitation bath has been precipitated longer. The concentration profiles of the precipitant in the polymer film are shown schematically in Figure 1. Because of the flat concentration profiles of the precipitant over the cross section of the casting solution. the casting solution divides into a polymer-rich and a polymer-poor phase. Formation of Symmetric and Asymmetric Membrane Structures. the precipitant. Therefore. and has a composition similar to the original casting solution A. and the solvent during the precipitation process are considered. a randomly distributed polymer structure is obtained during precipitation. Little solvent has diffused out of and little precipitant has diffused into the layer. the precipitant is added to the casting solution by immersing the cast polymer film in a bath of the precipi-
.30
Handbook of Industrial Membrane Technology
During precipitation. or.22 for various times. the precipitation is slow. its viscosity is therefore higher. or syneresis. the concentration profiles in the cast film are flat. of the solid polymer accompanying this composition change produces stresses in the polymer. During this time. the solid. these stresses cannot be as easily relieved by bulk movement of polymer as in the fluid polymer layer.6. Thus. the polymer structure either slowly undergoes creep to relieve the stress. The significant feature of the vapor-phase precipitation process is the slow diffusion of precipitant from the vapor phase adjacent to the film surface into the polymer solution. the viscosity of the precipitated polymer climbs from the point of first separation until it becomes almost solid at the point of gelation. has lost solvent and has gained precipitant. (3) The solid polymer layer: In this layer. In the second membrane precipitation procedure. In this layer. Two different techniques have been employed for the precipitation of membranes from a polymer casting solution. Shrinkage. if the stress builds up too rapidly to be dissipated by creep.

The most important features in immersion precipitation are the steep concentration and activity gradients of all components found in the polymer solution close to the polymer-precipitation medium interface.
I membrane I I
structure
at t = 4
’ I I I I I
point
of solidification
point
of precipitation
water
vapor
poly met.23.film
the concentration profiles times during the formation of the preof a sym-
Figure 1.
. These profiles over the cross section of the cast polymer film are shown schematically in Figure 1.Synthetic
Membranes
and Their
Preparation
31
tation fluid. This structure and formation can again be understood by considering the concentration profiles of the polymer. the precipitation is rapid. and an asymmetric membrane structure is obtained. the solvent and the precipitant during the precipitation process. In this case.22: Schematic diagram showing cipitant in the casting solution at various metric structure membrane.

32
Handbook of Industrial Membrane Technology
I-
precipitant
-
I
I membrane structure at t = 4
I
.
When the cast polymer film is immersed into the precipitation bath.23: Schematic diagram showing concentration profiles of the precipitant in the casting solution of various times during the formation of an asymmetric skin-type membrane. where. however. Phase separation therefore occurs initially at the surface of the film. the precipitant concentration is still far below the limiting concentration for phase separation. In the interior. solvent leaves and precipitant enters the film. the concentration of the precipitant soon reaches a value resulting in phase separation. there is a net movement of the polymer perpendicu-
.
I
water
bath
polymer
film
Figure 1. due to the very steep gradient of the chemical potential of the polymer. At the film surface.

Depending on the resistance of the skin to the flux of precipitant. Uniform and Graded Pore Structure of Skin-Type Membranes. A uniform pore structure is obtained when the concentration profiles are flat and the time that the system needs to move from the point of precipitation to the point of solidification is about the same over the entire film cross section.Synthetic Membranes and Their Preparation
33
lar to the surface. the exchange of solvent and precipitant is much faster than through the unfractured skin.
. such as syneresis. However. and the concentration profiles in the casting solution interior become less steep. The formation of the skin is identical with that of the sponge-structured membranes. shrinkage of the polymer causes it to propagate by draining the freshly precipitated polymer at the bottom of the finger to the side of the finger. the same situation in the sublayer is obtained as in a membrane precipitated from the vapor phase. also play an important role. With finger-structured membranes. Thus. This skin also serves to hinder further transport of precipitant into and solvent out of the casting solution. One is a sponge-like structure and the other is a finger-like substructure underneath the skin. The points at which the skin has been fractured form the initiation points for the growth of the fingers. and a structure with uniform randomly distributed pores is formed. and the precipitation front moves much faster within a finger than in the casting solution bypassed between fingers. shrinkage. leading to a more or less flat concentration profile of the precipitant in the casting solution just beneath the skin. The formation of finger-structured membranes is conveniently divided into two steps: the initiation and the propagation of fingers. skin-type membranes. The skin thus becomes the rate-limiting barrier for precipitant transport into the casting solution.43 Other phenomena. the two characteristic structures shown in Figure 1 . A graded pore distribution is obtained when the time between precipitation and solidification of the polymer increases with increasing distance from the skin. shrinkage stress in the solid polymer skin cannot be relieved by creep relaxation of the polymer and the homogeneous layer ruptures. It is the concentrated surface layer which forms the skin of the membrane. Within a finger. once the precipitated skin is formed. In spongestructured.and ‘Finger”-Like Structures.24 which shows the growth of a finger at various times. as a result of syneresis. and stress relaxation in the precipitated polymer. Once a finger has been initiated. Skin Type Membranes With “Sponge’. the concentration profile may vary from a completely flat one (virtually no concentration gradient over the cross section of the cast polymer film) to a concentration profile showing an initially steep gradient at the beginning of precipitation which decreases as precipitation proceeds through the polymer film. the formation process is more complex and cannot entirely be described by the thermodynamic and kinetic arguments of phase separation processes. This leads to an increase of the polymer concentration in the surface layer. Precipitation therefore occurs much slower and a sponge-like structure is formed between the fingers.I4 are obtained. In skin-type membranes. This is schematically shown in Figure 1. The formation of the sponge-structured membranes can also be rationalized by the precipitation process described above. This solution is protected from immediate exposure to the precipitant by a layer of precipitated polymer. it is assumed that the diffusion of the precipitant through the skin is the rate-limiting step.

Several authors have commented on the two different structures shown in Figure 1.. i.24: precipitation.-
--
.--
4
jpolymer
solution
J
+skin
Figure 1. completely surrounding droplets of the polymer-rich phase..precipitant-
-
_-
-. the continuous phase is formed by the polymer-rich phase.34
Handbook
of Industrial
Membrane
Technology
skin 1
.
. The precipitant and the solvent used in membrane preparation determine both the activity coefficient of the polymer in the solvent-precipitant mixture and the concentration of polymer at the point of precipitation and solidification.----. before the solidification of the polymer-rich phase.e. The Effect of the Solvent-Precipitant System on Membrane Structure and Properties. It is generally assumed that a fine structure is already predetermined in the casting solution prior to the actual phase separation.
Schematic
diagram
of finger
formation
at various
times
during
Cellular and Modular Membrane Fine Structure.
--
1skin
r-
--jiTcij3tant _---.15.44r4s A cellular structure will be obtained when in the beginning state of the phase separation. A nodular structure is obtained when the precipitant-rich phase is continuous.

the original acetone-formamide ratio of the solvent system will be changed. 45% acetone and 30% formamide do not show any NaCIrejection unless they are annealed hot water. organic solvents. and tend to prevent initiation of fingers.35 As indicated earlier. due to a loss of solvent. 45% acetone.26 where the flux and salt rejections of membranes precipitated in ice water from a casting solution of 25% cellulose acetate.46 Meanwhile. while high polymer concentrations tend to induce precipitation of sponge-structured membranes. at longer evaporation times (>2 minutes). This is shown in Figure 1. or other “porefarmers” to the casting solution or precipitation bath is described in detail in the literature. it has become apparent that skin-type asymmetric membranes can be prepared without an evaporation step. Thus. The effect of polymer concentration on membrane structures can be explained by the initiation and propagation of fingers. A low polymer concentration in the casting solution tends to induce precipitation in a finger-type structure. Higher polymer concentration in the casting solution produces a higher polymer concentration at the point of precipitation. where the salt rejection of cellulose acetate membranes prepared from a casting solution containing 25% cellulose-acetate.25. Increasing the viscosity of the casting solution has the same effect. it was shown. The effect of the evaporation results in an increase of the polymer concentration at the surface of the cast film. which was exposed to air at room temperature prior to precipitation in water at O’C and annealing at 75°C for 2 minutes is shown as a function of the evaporation time. Thus. The Effect of Pre. that cellulose acetate membranes made from casting solutions with formamide concentrations exceeding 35% did not reject NaCl. which will thus tend to increase the strength of the surface of polymer film first precipitated. In an independent set of experiments. the evaporation step prior to the precipitation of the film was considered as being essential for the formation of the asymmetric membrane. for this type of membrane the annealing is an essential postprecipitation membrane treatment step.and Post-Precipitation Procedures on Membrane Structures and Properties.
. the solvent or the precipitant and with that the rate of precipitation. acetone will preferentially be evaporated because it is the more volatile solvent. This is due to a significant change in the casting solution composition at the membrane surface because of differences in the boiling temperatures of acetone (kp = 6O’C) and formamide (kp = 12O’C). At short evaporation times (<I minute). 45% acetone and 30% formamide are shown as a function of the annealing temperature. The effect of the annealing procedure on membrane flux and sale rejection is demonstrated in Figure 1. The filtration tests were carried out at a hydrostatic pressure of 100 bar with a 1% NaCl solution. During the air exposure time. The annealing time was kept constant at 2 minutes and the membranes were tested at 100 bars with a 1% NaCI-solution. there is a drastic loss in rejection. The effect of additives such as salts.Synthetic Membranes and Their Preparation
35
The effect of additives to the casting solution or to the precipitant on the membrane structure can be explained by changes of the activity coefficients of the polymer. In the original recipes for making phase inversion membranes. while the formamide concentration will increase in the film surface. and 30% formamide.36p47 The Effect of the Polymer Concentration in the Casting Solution on the Membrane Structure. there is no change in rejection. however. membranes made from a casting solution containing 25% cellulose acetate. This does not necessarily lead to denser membranes.

45% acetone and 30% formamide by precipitation in water as a function of the past precipitation annealing temperature.
-1 100
s3-G C 0 c
v
al
2
L
50
i
I
I
I
I
\
. (Test conditions: 1% NaCI-solution.
20 Annealing
40
60 temperature
80 (OC)
100
Figure 1. 45% acetone and 30% formamide by precipitation in a water bath at 0°C as a function of the evaporation time prior to the precipitation.26: Transmembrane flux and salt rejection of a membrane prepared from a solution of 25% cellulose acetate.
.36
Handbook of Industrial Membrane Technology
I
I
I
3
I 4
I
5
1
2
Evaporation
time
(min)
Figure 1. (Test condition: 1% NaCI-solution. 100 bar hydrostatic pressure). 100 bar hydrostatic pressure).25: Rejection of a cellulose acetate membrane prepared from a casting solution containing 25% cellulose acetate.

For the separation of gases silicon rubber. Modern polymer chemistry is highly proficient in tailoring polymers to specific aims in terms of mechanical or thermal stability and chemical compatibility. nickel. cellulose esters and various polyamides serve as the barrier polymer for the membrane preparation. Homogeneous Polymer Membranes. The separation of various components in a solution is directly related to their transport rates within the membrane phase. Homogeneous Metal and Glass Membranes. physical properties and. which is determined mainly by their diffusivity and concentration in the An important property of homogeneous membranes is membrane matrix. There is only one type of homogeneous metal membrane of technical importance. etc. consistent with the required strength and absence of pinholes and defects. Because of their high selectivity. The mass transport in homogeneous membranes occurs strictly by diffusion.g4 The permeability of hydrogen in palladium alloy membranes is highly temperature dependent.55 Thus. such as glass or certain metals. Although there are a number of homogeneous membranes made from inorganic materials. which in general involve the separation of different low molecular weight components with identical or nearly identical molecular dimensions. However. pervaporation.64 The membranes generally consist of 10 to 50 pm thick metal foils. is the more widely used basic material. and the final product will represent a compromise between necessary strength.60r61 For reverse osmosis. in palladium.53r54. selectivity and mass-transfer rates. The most important applications of homogeneous polymer membranes are in gas separation. Homogeneous membranes should. palladium alloys and several other metals higher such as platinum. be as thin as possible. This is the palladium. is several orders of magnitude
than of any other gases. these
.Synthetic Homogeneous Membranes
Membranes and Their Preparation
37
A homogeneous membrane is merely a dense film through which a mixture of chemical species is transported under the driving force of a pressure. because of its relatively high permeability. thus permeabilities are rather low.
The permeability of hydrogen silver. s6 The principle aim is to create as thin a barrier as possible. therefore. and reverse osmosis. crystallization and orientation are to be avoided as much as possible when high permeabilities and transmembrane fluxes are desired. therefore. concentration. The membrane phase itself may be solid or liquid. the mechanical strength of the polymer as well as its selectivity may then be adversely affected. iron. or palladium-alloy membrane used for the separation and purification of hydrogen. by extruding from a polymer melt or by blow and press molding. and hence similar diffusivities.Flat sheets can be prepared by casting from solution. Hollow fibers are generally made by extrusion with central gas injection. The two basic membrane configurations are flat sheets and hollow fibers.5~5*.4*42 that chemical species of similar size. The separation is. differs signigicantly.59 Because of their high selectivity for different chemical components homogeneous membranes are used in various applications. may be separated when their concentration. that is their solubility in the film. carried out at elevated temperature (m4000C).the technically more important structures are of polymeric origin. mass transfer will be greater in amorphous polymers than in highly crystalline or cross-linked polymers.62r63 Most technically utilized homogeneous polymer membranes consist of a composite structure where a very thin homogeneous selective polymer film is supported by a thicker microporous structure providing the mechanical strength. or electrical potential gradient. in particular. In general.

and therefore about 100 times the thickness of the selective barrier of an asymmetric polymer membrane. Otherwise.1n this configuration.n Unsupported Liquid Membranes.mt71. Very thin unsupported liquid membranes may be obtained. a pore size small enough to support the liquid membrane phase sufficiently under hydrostatic pressure and the polymer of the substructure should be hydrophobic in nature for most liquid membranes used in contact with aqueous feed solutions. In order to avoid a break-up of the film. when the selective membrane material is stabilized by an aPPro_
. a microporous polymer structure is filled with the liquid membrane phase.e. such as a Goretex (Gore Corp. when used in aqueous solutions.) type stretched polytetrafluorethylene or polyethylene membrane.65 The same is true for the use of homogeneous silica glass membranes. the microporous structure provides the mechanical strength and the liquid-filled pores the selective separation barrier. thus they are used as the selective barrier in pH-electrodes. there are only very few commercial plants in operation. the selective liquid barrier material is stabilized as a thin film by a surfactant in an emulsion-type mixture. the only other homogeneous inorganic material which shows any promise to be used as selective barrier especially for the separation of helium.n Supported Liquid Membranes The preparation of supported liquid membranes is extremely simple. Homogeneous glass membranes also have a high selectivity for H+-ions. The microporous substructure should have a high porosity. the useful lifetime of the membrane is rather limited. Liquid membranes have gained increasing significance in recent years in combination with the so-called facilitated transport which utilizes selective “carriers” transporting certain components such as metal-ions selectively and at a relatively high rate across the liquid membrane interphase. however. no commercial industrial size plants are in operation. The disadvantage of supported membranes is their thickness which is determined by the thickness of the microporous support structure. high boiling point and. In the first case.. Both types of membranes are used today on a pilot-plant stage for the selective removal of heavy metalions or certain organic solvents from industrial waste streams. i. It is difficult. In practice. Liquid Membranes. which is in the range of IO to 50 pm. Until today. Two different techniques are used today for the preparation of liquid membranes. Like metal membranes.38
Handbook of Industrial Membrane Technology
membranes are used for production of high purity hydrogen (>99. Their preparation is described in some detail in the literature and shall not further be discussed in this chapter.6ar6g . in the hydrophobic liquid which may consist of a selective carrier such as certain oximes or tertiary or quaternary amines dissolved in kerosene. Thus. when certain requirements concerning the selective barrier and the microporous support material are fulfilled. Although the process seems technically feasible.99% Hz). to maintain and control this film and its properties during a mass separation process. however. liquid membranes are prepared by soaking a hydrophobic microporous membrane.) or Cellgard (Celanese Corp. They have also been used rather effectively for the separation of oxygen and nitrogen. The liquid membrane material should have a low viscosity and low vapor pressure. glass membranes are operated at elevated temperature. some type of reinforcement is necessary to support such a weak membrane structure.In the second technique for making liquid membranes. a low water solubility. the fluxes of supported liquid membranes can be low even when their permeabilities are high.66t67 It is relatively easy to form a thin fluid film.

27.
stripping
agent
\
liquid
membrane
mode. most of them based on zeolites and bentonites. the hydrophobic membrane phase and the surrounding aqueous phases are more fractionated and the diffusion pathways become longer as a result. Although there are a number of inorganic ion-exchange materials. The preparation procedure is indicated in Figure 1.
trtatt’“tnt
practical
application
\Ccparation
liquid mechanism
mcmbtain unsupported liquid mrmbranrs
Figure 1. there are many possible types with different polymer matrixes and different functional groups to confer ion-exchange properties on the product. With another aqueous solution.
. the component to be eliminated is supplied to the original emulsion and passes through the membrane into the internal solution. in which the aqueous phase is surrounded by a relatively thin hydrophobic membrane forming phase which is surrounded by a second aqueous phase.Synthetic Membranes and Their Preparation
39
priate surfactant in an aqueous emulsion. droplets form in this process. Ion-exchange membranes consist of highly swollen gels carrying fixed positive or negative charges. Ideally. A hydrophobic membrane phase is transformed into an emulsion by stirring with an aqueous phase. The mass exchange occurs between the inner and outer aqueous phases through the liquid membrane interphase. The properties and preparation procedures of ion exchange membranes are closely related to those of ion-exchange resins. In reality.
Schematic diagram showing the formation of an emulsified liquid
Ion-Exchange Membranes. these materials are rather unimportant in ion-exchange membranes and will not be discussed further.74 As with resins.27: membrane.

they exclude all cations and are permeable to anions only. the positive counterions. In contrast. and the negative co-ions. and (2) anion-exchange membranes which contain positively charged groups fixed to the polymer matrix. the fixed anions are in electrical equilibrium with mobile cations in the interstices of the polymer.
l
0
l
&Matrix
with ounterCo-Ion Ion
Fixed
Charges
OC
0
Figure 1.40
Handbook of Industrial Membrane Technology
There are two different types of ion-exchange membranes: (1) cation-exchange membranes which contain negatively charged groups fixed to the polymer matrix.
. Due to the exclusion of the co-ions. Therefore. The most required properties from ion-exchange membranes are:
l
High permselectivity-an ion-exchange membrane should be highly permeable for counter-ions. High chemical stability-the membranes should be stable over a pH-range from 1 to 14 and in the presence of oxidizing agents. a cation-exchange membrane permits transfer of cations only. In a cation-exchange membrane. but should be impermeable to co-ions.28. are more or less completely excluded from the polymer matrix because of their electrical charge which is identical to that of the fixed ions. This figure shows schematically the matrix of a cation-exchange membrane with fixed anions and mobile cations. Anion-exchange membranes carry positive charges fixed on the polymer matrix. as indicated in Figure 1. which are referred to as counter-ions. the mobile anions. called co-ions.28: Schematic diagram of the structure of a cation-exchange membrane showing the polymer matrix with the negative fixed charges. Low electrical resistance-the permeability of an ion-exchange membrane for the counter-ions under the driving force of an electrical potential gradient should be as high as possible. Good mechanical and form stability-the membrane should be mechanically strong and should have a low degreeof swellingor shrinking in transition from dilute to concentrated ionic solutions.

the most important of which are relatively high electrical resistance and poor mechanical strength when highly swollen in dilute salt solutions.. Accordingly.g. $ N+ . is completely dissociated over nearly the entire pH-range. and the chemical and thermal stability. Therefore.*l -NHJ+ . The properties of ion-exchange membranes are determined by two parameters. since the fixed ion charges are distributed homogeneously over the entire matrix. :NH2+ . while the carboxylic acid group -COO-is virtually undissociated in the pH-range <7. according to their structure and preparation procedure... 5s’. The following moieties are used as fixed charges in cation-exchange membranes:78-84
-SOB.e. homogeneous and heterogeneous membranes. is affected by the fixed charge concentration. The basic polymer matrix determines to a large extent the mechanical. causes a high degree of swelling combined with poor mechanical stability.
-coo. chemical and thermal stability of the membrane.
-Po32-
. in general. For instance. i. but also has a large effect on the electric resistance and the permselectivity of the membrane.s2 Most commercial ion-exchange membranes can be divided.. In general. Very often the matrix of an ion-exchange membrane consists of hydrophobic polymers such as polystyrene.e. e.76r77. fixed charges may be:80. especially. The degree of swelling..75 Although these basic polymers are insoluble in water and show a low degree of swelling. but they also have a significant effect on the mechanical properties of the membrane. The type and the concentration of the fixed ionic charges determine the permselectivity and the electrical resistance of the membrane. The degree of crosslinking then determines to a large extent the degree of swelling.
These differently charged groups have a significant effect on the ion-exchange behavior of the membrane. The quaternary ammonium group again is completely dissociated over the entire pH-range. they may become water soluble by the introduction of the ionic moieties. the basic polymer matrix and the type and concentration of the fixed ionic moiety. The methods of making homogeneous ion-exchange membranes can be summarized by three different basic procedures:
. The sulfonic acid groups. -SO.Synthetic Membranes and Their Preparation
41
It is often difficult to optimize the properties of ion-exchange membranes because the parameters determining the different properties often act contrary. the polymer matrix of ion-exchange membranes is very often crosslinked. while the primary ammonium group is only weakly dissociated. Homogeneous ion-exchange membranes have significantly better properties in this respect. polyethylene or polysulfone. a high degree of crosslinking improves the mechanical strength of the membrane but also increases its electrical resistance. Preparation Procedure of Homogeneous Ion-Exchange Membranes. ion-exchange membranes are referred to as being weak or strong acid or basic in character. heterogeneous ion-exchange membranes have several disadvantages. -AsOj2-
In anion-exchange membranes. A high concentration of fixed ionic charges in the membrane matrix leads to a low electric resistance but. i. into two major categories.

.-&~2-~r~~. polymerizing the imbibed monomer... The cation-exchange membrane is obtained by the following reaction scheme:83
HC=CH2
.
803-H+ S 03. (2) Introduction of anionic or cationic moieties into a preformed film by techniques such as imbibing styrene into polymer films. . Starting with a film makes the membrane preparation rather easy. Excess monomers are removed by washing the film in water.. such as cellophane or polyvinyl alcohol. or a film from a hydrophobic polymer. The solution is cast into a film. The first membranes made by this procedure were prepared from phenol by polycondensation with formaldehyde according to the following reaction scheme:‘*
&%&.
. H+ 3-
Phenol is treated with concentrated Hz!504 at elevated temperatures leading to phenolsulfonic acid in paraform.or anion-exchange membrane is the polymerization of styrene and divinylbenzene and subsequent sulfonation or amination.
The anion-exchange group is introduced into the polymer by chloromethyiation and amination with trimethylamine according to the following reaction scheme:
There are numerous literature references to the preparation of ion-exchange membranes by polymerization. A very common method of preparing a cation.-CH-CHz
-. such as polyethylene
. a brown crystalline material. The starting material may be a film from a hydrophilic polymer.42
Handbook of Industrial Membrane Technology
(I) Polymerization or polycondensation of monomers... and then sulfonating the styrene. of which at least one must contain a moiety that either is or can be made anionic or cationic.....-CC-CHz-. The phenolsulfonic acid and a solution of formaldehyde in water is treated at elevated temperatures for several hours.

-
FH -CH.cHj +
SO.o
Amination
-$H-CH2-CY sop . followed by dissolving the polymer and casting it into a film. Membranes made by any of the above methods may be cast around screens or other reinforcing materials to improve their strength and dimensional stability. R W -7 -Cl+ + HCI
So2 -NH-&H3 R
Quaternization
+Cl$Br __. leading to a reinforced membrane
.” Sulfochlorination
-CH~CH.-CH.
-SrH-CH2-CH2 SO2 -NH-$+CH.
Anion-exchanog
-CH-CH2-CH3 I I FH3 SO2 -NH-+CH3 R Bi
memhmnp. SC&Cl + sNa OH __.Synthetic Membranes and Their Preparation
43
or polystyrene.-CH_j SW
l
HCI
Hydrolysis
-fH-CH. ion-exchange membranes made by sulfochlorination and amination of polyethylene sheets have low electrical resistance combined with high permselectivity and excellent mechanical strength. The reaction scheme for the sulfonation of polysulfone is as follows:*5
[-Q-
o-~-~. The reaction scheme for preparation of ion-exchange membranes by sulfochlorination and amination of polyethylene sheets is given below.*CI.~-O-o-
SO*-] +S03
+NaOH
-[-Q-o-~-~~-o-~-302-]
B'OiNa+
n
The sulfonated polysulfone can be cast into a film on a screen and precipitated after most of the solvent has been evaporated.+ I$ N+-CH. R
(3) Introduction of anionic or cationic moieties into a polymer chain such as polysulfone. -CH-C%CHi &O.Na+ + NaCl + H.

largely due to the risk of precipitation of inorganic or organic negatively charged materials. there are numerous methods reported for the preparation of ion-exchange membranes with special properties s’-~ for instance. Most heterogeneous membranes that possessadequate mechanical strength generally show poor electrochemical properties. and the polymerization subsequently completed. . and the solvent evaporated to give an ionexchange membrane. Another procedure is dry molding of inert film-forming polymers and ion-exchange particles and then milling the mold stock. -_( CF. These membranes consist of fine colloidal ionexchange particles embedded in an inert binder such as polyethylene. the limiting current density through the anion-exchange membrane is lower than in the cation-exchange membrane. ion-exchange particles can be dispersed in a solution containing a film-forming binder. To overcome this problem. Since these ion-exchange particles swell when immersed in water. Similarly.. for use as battery separators. phenolic resins or polyvinyl chloride. As a consequence.44
Handbook of Industrial Membrane Technology
with excellent chemical and mechanical stability and good electrochemical properties. In the literature. On the other hand. it has been difficult to achieve adequate mechanical strength combined with low electrical resistance. They combine good chemical stability with high selectivity and low electric resistance.Na
COONa
Significant effort has also been concentrated on the development of anionexchange membranes with low fouling tendencies. In a conventional electrodialysis plant.-.8
m/n "' as-40 al-4
SO. In
. or in the’chlor-alkali process. The structure of a membrane produced by Asahi Chemical is shown in the following diagram:W r CF?-CF2). Special Property Membranes. its electrical resistance may rise during operation. Also.F 1” I! & 0 !F
I *
cFJ:F ?
(
y24
iz CFJFF
0 I
( fVq
l/tm t n) 1 6 . which are characterized by the fact that they can be used at a higher current density. a membrane that contains ion-exchange particles large enough to show an adequate electrochemical performance exhibits poor mechanical strength. ion-selective electrodes. Heterogeneous membranes with usefully low electrical resistances contain more than 65% by weight of the crosslinked ion-exchange particles. Heterogeneous Membranes.66 Such a membrane can be prepared simply by calendering ion-exchange particles into an inert plastic film.. Especially membranes recently developed for the chloralkali industry are of commercial significance. different companies produce special anionic membranes.-CF I.r CF. ion-exchange particles can be dispersed in a partially polymerized binder polymer. These membranes are based on polytetrafluoroethylene and carry sulfone groups in the bulk of the membrane phase and carboxyl-groups on the surface as the charged moiety.

many polymers with satisfactory selectivities and permeabilities for the various components in gas mixtures or liquid solutions are not well suited for the preparation of asymmetric membranes by the phase inversion process described before. In an integral asymmetric membrane. A composite membrane is shown schematically in Figure 1. as indicated before. The performance of a composite membrane is not only determined by the proper-
. and which are therefore not suited for preparation into integral asymmetric membranes.
/Porous
Support
Figure 1. In making composite membranes. the actual mass separation is achieved by a solution-diffusion mechanism in a homogeneous polymer layer. This preparation mode leads to significant advantages of the composite membrane compared to the integral asymmetric membrane. One is the preparation of a suitable microporous support and the second task is the preparation of the actual barrier layer and laminating it to the surface of the support film. Composite membranes differ from asymmetric reverse osmosis membranes by the mode of preparation which consists of two steps: (I) casting of the microporous support. Composite Membranes In processes such as reverse osmosis. Unfortunately. the selective barrier layer and the microporous support always consist of the same polymer. an asymmetric membrane structure is mandatory for these processes. two completely different tasks have to be solved.Synthetic Membranes and Their Preparation
45
general. Therefore. This means polymers which show the desired selectivity for a certain separation problem. This. gas separation and pervaporation. of course.tm thick microporous film. these membranes should be as thin as possible. In a composite membrane. may well be utilized as the selective barrier in composite membranes. and (2) deposition of the barrier layer on the surface of this microporous support layer.29: Schematic diagram of an asymmetric composite membrane showing the microporous support structure and the selective skin layer. Since the diffusion process in a homogeneous polymer matrix is relatively slow.29. different polymers may be-and in general are-used for the microporous support and the selective barrier layer. This has led to the development of the socalled composite membranes. but have poor mechanical strength or poor filmforming properties. the permselectivity of these membranes is lower than of regular membranes. Preparation Procedures of Composite Membranes. expands the variety of available materials for the preparation of semipermeable membranes. It is composed of a 20 to 100 nm thin dense polymer barrier layer formed over an approximately 100 I.

e. e. Interfacial polymerization of reactive monomers on the surface of the microporous support film. which are available as still soluble prepolymers that can easily be crosslinked by a heat treatment procedure thus becoming insoluble in most solvents. Today. the prepolymer is unable to penetrate the support and a rather thin uniform barrier layer of 0. Dip-coating of the microporous support film in a polymer. are suited for the preparation of this type of composite membrane. The techniques used for the preparation of composite structures may be grouped into four general types:91+5 (I) (2) Casting of the barrier layer separately.97 Particularly.46
Handbook of Industrial Membrane Technology
ties of the selective barrier layer.% Dip coating a microporous support membrane in polymer or prepolymer solution was also first developed for the preparation of reverse osmosis membranes. Casting an ultrathin film from a dilute acetone solution on a glass plate and releasing the film from the plate after the evaporation of the acetone by immersion in water was another method of preparing ultrathin selective barriers. Gas-phase deposition of the barrier layer of the microporous support film from a glow discharge plasma.. a reactive monomer or prepolymer solution followed by drying or curing with heat or radiation. which consisted. The actual selective barrier was prepared by dissolving 2-10 percent cellulose diacetate in a solvent exhibiting slight water solubility such as cyclohexanone. which had better overall mechanical and thermal stability and which was insoluble in the solvent of the barrier layer. from dissolving the support membrane. such as an “open” polysulfone ultrafiltration membrane. a microporous membrane prepared from mixed cellulose esters was first coated by a protective layer of polyacrylic acid to prevent the solvent of the casting solution of the barrier layer.g.g. Although both preparation techniques lead to barrier layers of less than 100 nm with correspondingly high flux rates. polymers such as polydimethylsiloxane. but it is also significantly effected by properties of the microporous support film.30.05 to 1 pm thickness can easily be prepared.
(3) (4)
Casting an ultrathin film of cellulose acetate on a water surface and transferring the film on a microporous support was one of the earliest techniques used for preparing composite reverse osmosis membranes for water desalination. of cellulose triacetate in chloroform. they are not well suited for large scale industrial production. dip coating is applied mainly for the preparation of composite membranes to be used in gas separation and pervaporation.
. A typical composite membrane prepared by dip coating an asymmetric polysulfone ultrafiltration membrane into a 1 wt % solution of polydimethylsiloxane followed by thermal crosslinking is shown in the scanning electron micrograph of Figure 1. on the surface of a water bath followed by lamination to the microporous support film. preparation and Deposition of the Selective Barrier Layer on the Microporous Support.. This technique was later improved by using a microporous sublayer. If the pore dimension in the support membrane is selected properly. Here.

by far the most important technique for preparing composite membranes is the interfacial polymerization of reactive monomers on the surface of a
. 98 large scale industrial production utilizing plasma polymerization for the preparation of composite membranes seems to be difficult. and fluxes of 1.
electron micrograph showing a composite membrane as the selective layer deposited on a polysulfone sup-
Gas phase deposition of the barrier layer on a dry microporous support membrane by plasma polymerization was also successfully used for the preparation of reverse osmosis membranes. The plasma reactions are rather heterogeneous not only involving polymerization but depolymerization and modification of functional groups.30: Scanning with polydimethylsiloxane port membrane. Today.5 m3mzd-’ when tested with seawater have been prepared on a laboratory scale. Although reverse osmosis membranes with excellent desalination properties showing salt rejection in excess of 99%. Many organic compounds having adequate vapor pressure can be used to form a barrier layer on a microporous support.Synthetic
Membranes
and Their
Preparation
47
Figure 1.

31: Schematic diagram showing the formation brane by interfacial polymerization of polyethyleneimine anate. The mechanical stability of the membrane as well as its flux rate is to a large extent determined by the pore size and the overall porosity of the substructure as can easily be demonstrated by the schematic diagram of Figure 1. its overall performance is strongly affected by the microporous substructure. Al though the selectivity of a composite membrane is determined nearly exclusively by the actual barrier layer. which exhibited water fluxes of about 1 m3m3d-’ and salt rejections in excess of 99% when tested with seawater at 60 bar pressure. which was reacted interfacially at the membrane surface with a 0. This diagram shows an idealized homogeneous barrier layer with a
. the polyethyleneimine reacts rapidly at the interphase with the toluene diisocyanate to form a polyamide surface skin while amine groups below this surface remain unreacted. produced by Film Tee Corporation from monomeric diamine reactants such as m-phenylenediamine interfacially polymerized with trimesoyl chloride shows not only excellent desalination properties but also highly improved stability towards oxidizing agents. the FT-30.31. A polysulfone support membrane was soaked in an aqueous solution of 0. The first membrane produced on a large scale with excellent reverse osmosis desalination properties was developed in the early seventies in the North Star Research Institute under the code name NS 100.ree Preparation of the Microporous Support of Composite Membranes.48
Handbook of Industrial Membrane Technology
microporous support film. In the second heat treatment step. In a first step. internal crosslinking of the polyethyleneimine takes place. (2) a thin crosslinked polyethyleneimine layer which extends into the pore of the support film.” The preparation procedure of this membrane.2 to 1% solution of toluene diisocyanate in hexane. A heat curing step at 110°C leads to further crosslinking of the polyethyleneimine. Thus. was rather simple. further improvements were achieved by using aromatic diamines and triacyl chloride reactants.32.
of a composite memwith toluene diisocy-
Although the NS-100 membrane showed significantly higher salt rejection capability and higher fluxes than most integral asymmetric reverse osmosis membranes. The process seems to involve two types of reactions. and (3) the actual polysulfone support membrane.
PEI
+
Figure 1. the final membrane has three distinct layers of increasing porosity: (I) the dense polyamide surface skin which acts as the actual selective barrier.5 to 1% polyethyleneimine. The preparation is shown schematically in Figure 1. One of these membranes.

r. Unfortunately. For mechanical strength.
M
Figure 1. hM. to a first approximation..32: Schematic diagram showing a homogeneous barrier film on a microporous substructure.1 pm on a microporous support. ho.e.M
= E X0 + (l--E)
This relation which can be obtained by simple geometric considerations indicates that the effective diffusional path length is always longer than the thickness of the barrier layer and that it is strongly affected by the overall porosity of the substructure. For practical purposes. in excess of 50%.2 pm and a porosity of 50%. and the radius of the pores. The actual diffusion pathway of thecomponent through the barrier layer is always longer than the thickness of the layer. the thickness of the barrier layer. in the support structure. the effective diffusional length is approximately identical with the thickness of the barrier layer it will increase according to equation (12) by about an order of magnitude. E. i. be expressed as a function of the membrane overall porosity. the porosity of the support should be as high as possible. many ultrafiltration membranes have a rather low surface porosity of 24%. The geometrical consideration expressed in equation (12) indicates that the surface porosity of the support layer will significantly effect the membrane flux. Accordingly. While at relatively high surface porosities. the flux will decrease by the same magnitude.Synthetic Membranes and Their Preparation
49
thickness of 0. with the average pore diameter being 0. Industrial Scale Membrane Production For any membrane to be useful in an industrial process it has to be produced on a large scale and installed into an appropriate device which should be com-
. and it should always be as high as possible when optimal flux rates shall be achieved with thin film composite membranes. by the following relation: (12)
L. The average path length. when the porosity is less than 1% assuming the barrier layer thickness is approximately the same as the pore radius. the pore diameter should not be significantly larger than the film thickness. The actual average diffusional pathlength hi of a component through a film of the thickness ho is expressed as a function of the overall porosity and the pore radius. can.

The pleated cartridge filter module. when asymmetrically structured. only six basic types are used today on a large industrial scale. which. It consists of a pleated membrane cartridge installed in a pressurized housing. porous membrane support plates. as indicated in Figure 1. and spacers forming the feed flow channel are clamped together and stacked between two endplates as indicated in the schematic diagram of Figure 1. is collected in the shell tube.“’ A variation of the basic plate-and-frame concept is the spiral-wound module. Their useful life is limited due to plugging of the membrane pores by retained solutes. The pressurized feed solution flows down the tube bore and the permeate is collected on the outer side of the porous support pipe. reliable. Membrane Modules and Their Fabrication. which is shown schematically in Figure 1. A large number of different module systems are described in the literature. chemical engineering aspects are of prime importance for the design of membrane modules. These modules are shown schematically in Figure 1. Cartridge type filters are operated at relatively low hydrostatic pressures.33 (e). The capillary membrane module requires as basic material membranes in a self-supporting capillary configuration.33 (c). While the previously described three membrane modules required flat sheet membrane material for their preparation.50
Handbook of industrial Membrane Technology
pact. The actual cartridge is made by pleating a membrane sheet and potting the ends by an appropriate resin or hot-melt-glue as indicated in Figure 1. The feed flow channel spacer. the membranes are cast directly on the porous pipes and in others they are prepared separately as tubes and then installed into the support pipes.5 cm. However. only slight variations in their basic configuration. There are various types of plate-and-frame modules on the market which offer. therefore. The feed solution enters the filter from the housing side and the product is collected in a center tube which is sealed against the housing by an O-ring. special membrane configurations are needed for the preparation of the tubular. and gas separation. consists of a large number of membrane capillaries with an inner diameter of 0. The diameters of tubular membranes are typically between l-2. In some modules.2 to 3 mm arranged in parallel as a bundle in a shell tube.33 (a). reverse osmosis. which permeates the capillary wall. which is shown in (a) is used mainly in dead-end microfiltration. carry the selective -/
. The tubular membrane module consists of membrane tubes placed into porous stainless steel or fiber glass reinforced plastic pipes. For technical use membranes are. Another module type used on an industrial scale for various membrane separation processes including ultrafiltration. The membranes. and the porous membrane support are rolled up and inserted into an outer tubular pressure shell.33 (a)-(f). capillary. The capillary membrane module. the membrane. Besides economic considerations. the membrane modules used for the various processes are equally different. ultrafiltration. The filtrate is collected in a tube in the center of the roll.33 (d). and gas separation is the plate-and-frame module.33 (b). and hollow fiber modules. however. Its basic design is illustrated in Figure 1. integrated into so-called modules. Its design has its origins in the conventional filter press concept. Since the various membrane separation processes differ significantly in their operational concept and their applications. The feed solution is passed down the center of the membrane capillary and the filtrate. and inexpensive. which is widely used today in reverse osmosis.

which have an outer diameter of 50 to 100 . tubular. In hollow fiber membranes.
.33 If).urn.Synthetic
Membranes and Their Preparation
53
barrier on the inner side of the capillary as indicated in the scanning electron micrograph of Figure 1. the membrane is collected as a flat sheet on the take up roll. which are obtained as up to 2 111 wide continuous sheets. A polymer solution is cast by a casting knife on a polyester or polyethylene support paper. which are installed as a bundle of several thousand fibers in a half loop with the free ends potted with an epoxy resin in a pressure tube as indicated in Figure 1. The filtrate passes through the fiber walls and flows up the bore to the open end of the fibers at the epoxy head.34. which shows a typical capillary ultrafiltration membrane prepared in a wet-spinning process. Flat sheet membranes are generally manufactured on a casting machine.
~
Figure 1. The cast polymer film is fed to the precipitation bath.
The same basic spinning process is used for the preparation of hollow fiber membranes. there are four basic membrane configurations produced today on a large scale. where residue solvent is removed. The feed solution is introduced around th'e outside of the hollow fibers. the selective layer is on the outside of the fibers. capillary. The membranes. which is shown schematically in Figure 1. which is continuously supplied from a roll. These are flat sheet. and hollow fiber membranes.34: SEM of a capillary
20 ~
IJm
membrane. are then further processed into the desired module configuration.35. After a certain residence time in a rinse bath. where the actual membrane is formed. Membrane Manufacturing Equipment
Based on the different modules used in technical scale membrane separation processes.

There are many variations of the basic membrane preparation techniques described in this chapter. the precipitation progresses from the inside to the outside and a capillary with the selective layer on the inside is formed immediately at the outlet of the spinneret. Since the outer diameter of the casting cone is slightly smaller than the inner diameter of the tube. thus. Hollow fibers have. The casting solution is always fed in the outer bore of the nozzle. the precipitant in the inner bore is replaced by an inert gas and the fiber is spun into the precipitation bath. have about the dimensions of the spinning nozzle. therefore. which is then converted into a membrane by immersing the entire tube into a precipitation bath. Between the precipitation bath and the spinneret there is an air gap as indicated in Figure 1.
Capillary and hollow fiber membranes are manufactured generally by a wetspinning process using a spinneret with a double bore nozzle as indicated in Figure 1. often significantly smaller diameters than the nozzle. therefore.36 (a).36: Schematic diagram indicating the function of a casting machine used for the preparation of supported flat sheet membranes. a thin polymer solution film is formed on the inside of the tube. If an asymmetric capillary membrane with the skin on the inside shall be produced.36 (b) where the fiber may be drawn to obtain the desired dimensions before precipitation.54
Handbook of Industrial Membrane Technology Tensioning Roller abric Roll Roller
Solution Trough \ Doctor
r-7
/Spreader
Tap Water Figure 1. If an asymmetric hollow fiber with the skin on the outside is to be produced. But most of them have only limited technical significance and shall not be discussed further. which shows a schematic diagram of a spinneret for the production of asymmetric hollow fiber or capillary membranes. the precipitant is directed through the central bore. The capillaries.
. Tubular membranes are prepared either by an ultrasonic welding process from flat sheet membranes or by direct casting on a porous support tube using a conical casting device pulled through the porous tube.

Steinkopff Verlag. Strathmann.tm. ultrafiltration. Handbook of Separation Techniques for Chemical Engineers. 6. M. research is concentrated on membranes having an asymmetric structure with high surface porosities in excess of 70% and average pore sizes at the surface between 0.
. they are all far behind the actual biological membranes. even for comparatively mature processes such as microfiltration and reverse osmosis. The main effort as far as synthetic membranes are concerned isconcentrated on the development of completely new membranes for processes such as pervaporation. lo2 Biocatalytic membranes. 4..56
Handbook of industrial Membrane Technology DEVELOPMENTS
FUTURE
For many separation processes. reverse osmosis. Controlled Release of Bioactive Materials. A. H.
REFERENCES 1. “better” membranes are desirable.These membranes are made from polymers. fuel cells or electrochemical production processes. gas separation.K.01 and 0. and electrodialysis in biotechnology and the chemical process industry..C. Mass..A. 3 (1985) 112-118. But. Sci. such as Nylon 6.1 E. Katchalsky. such as sea and brackish water desalination by electrodialysis and reverse osmosis or the production of ultrapure water by microfiltration.. (Ed. Porter. Schweitzer (Ed. Nonequilibrium Thermodynamics in Biophysics. Membrane filtration. Trennung von molekularen Mischungen mit Hilfe synthetischer Membranen. Baker. P. Strathmann. In reverse osmosis water desalination work is concentrated mainly on the development of membranes with improved stability against oxidizing agents such as chlorine or hypochlorite with higher fluxes and better flux stability. membranes with quite satisfactory properties are commercially available today. Cambridge.. 10 (1982) 81-181. 5. With the emerging use of microfiltration. or as ion transferring separators in batteries. pp.J. completely new requirements are put on the membranes to be used in these applications. Harvard University Press.T. A very important area of today’s ongoing research is the development of functional synthetic membranes which mimic the function of the biological membrane. in: P. Trendsin Biotechnology. 2. Darmstadt (1979). (1967). New York (1980). Academic Press. In microfiltration. and Curran.).. 2-3 to 2-103. energy and information transducing membranes have already been produced on a laboratory scalelo and are used today as biosensors for monitoring devices.). Liquid membranes with selective carriers used today for the separation and concentration of heavy metal ions or certain organic compounds are being developed further to be used in gas separation. Lonsdale.6 which show good solvent resistance and low adsorption for “fouling” materials. 3. Although a significant number of functional synthetic membranes have been developed. Memb. New York (1979). membrane distillation. H. McGraw-Hill. But basic research in membrane technology combined with progress in molecular engineering should help to improve the properties of functional synthetic membranes and thereby increase their use far beyond today’s level. H. R.

German water supplies were often devastated or contaminated by bombing raids. At the turn of the century (1906). It is astonishing that the art of controlling pore size and microstructure was developed to such a high degree of sophistication before understanding the mechanism of membrane formation. Conventional culturing techniques in liquid or gel-like nutrient media could take up 61
. Even today.” Commercial production began in 1927. researchers like Bigelow. A more efficient method for detecting coliform or pathogenic bacteria was needed. Porter
INTRODUCTION The beginnings of microfiltration (MF) can be traced back into the 19th century with the synthesis of nitrocellulose in 1845 by Schoenbein.2 Microfiltration
Mark C. Zsigmondy. Brown. Fick then used ether-alcohol solutions of the same (collodion) to form the first nitrocellulose membranes in 1855. Sartorius-Werke Aktiengesellschaft developed a commercial process for making cellulose nitrate membranes. In the first quarter of the twentieth century. Gertrude Mueller and others of the Hygiene Institute of the University of Hamburg developed the membrane for filtering and culturing bacteria. who was Director of the Institute of Colloid Chemistry of the University of Goettingen. but sales were largely confined to the laboratory market. Zsigmondy. Bechhold produced graded pore sizes in collodion membranes and measured the pore size with the “bubble-point” method (to be described later). The first important application of these membranes emerged during World War II. The membrane was known as the “Zsigmondy Membranfilter. the most common polymers used in MF membranes are mixed esters of cellulose-including cellulose nitrate. With the help of Professor Dr. and Bachmann made significant advances in the methodology of casting and regulating pore size. Gemberling. Schoep.

Each bacterium collected would grow into a colony of thousands overnight. The membrane could then be placed on top of a nutrient pad allowing the nutrients to diffuse up through the pores of the membrane to the bacteria on the surface of the membrane filter. Goetz had improved production methods to make membranes with higher flow rates and more uniform pore sizes. They are available as polycarbonate or polyester membranes. A unique feature of the “capillary-pore” membranes is that the pore size
. Millipore. By 1950.1). the colony would be readily visible and countable by the naked eye. MembranelEnka Gelman Millipore Nuclepore Norton Co. and Poretics Corp. the Lovell Chemical Company in Watertown. Gelman. to further develop the membrane. After the war (19471. and based on his findings.S. Gore Corp.. Eventually. Chemical Corps. The “capillary-pore” membranes are currently manufactured commercially only by Nuclepore Corp. Goetz visited Membranfiltergesellschaft (Sartorius). Alexander Goetz to Germany to obtain information on membrane filter production methods.62
Handbook of Industrial Membrane Technology
to 96 hours. Sartorius Co. the Lovell Chemical Company sold its membrane manufacturing facility to the newly organized Millipore Corporation. membranes made of materials other than cellulose nitrate began to appear: 1962 1963 1963 1963 1964 1970 1970 1975 1979 1980 1981 1984 Gelman Instrument Co. whereas the “tortuous-pore” structure resembles a sponge with a network of interconnecting tortuous pores. Joint Intelligence Objectives Agency sent Dr. was awarded a contract by the U. In 1954. S&S General Electric Selas Flotronics Celanese Co. Based on Goetz’s developments. Sartorius. the U. Ceraver Cellulose tri-acetate Regenerated cellulose Polyvinyl chloride and polyamide Polycarbonate Silver membrane Polypropylene Polytetrafluoroethylene Polypropylene Polysulfone Polyvinylidene fluoride Polyester Alumina
MEMBRANE
STRUCTURE
AND FABRICATION
All current MF membranes may be classified as either “tortuous-pore” or “capillary-pore” membranes (see Figure 2. The “tortuous-pore” membranes are the most common and include typical cellulosic membranes and virtually all other polymers.S. Massachusetts was awarded further contracts in 1952 to commercialize production. Mueller discovered that the entire bacterial flora in one liter of water could be deposited on a 47 mm diameter membrane within approximately 15 minutes. Other companies in both the United States and England then began to exploit the German technology base to manufacture membrane filters. In less that 24 hours of incubation (37”(Z). The “capillary-pore” structure is distinguished by its straight-through cylindrical capillaries. He also imprinted grid-lines on his filters to facilitate counting of bacterial colonies.

This is not possible with a "tortuous-pore" membrane since the pore openings do not correspond to the limiting pore size within the depth of the membrane.
.
Figure
2. The "tortuous pore" membrane appears to have some very large pore openings on the surface.1:
Capillary-pore
and tortuous-pore
membranes. This is demonstrated in Figure 2.Microfiltration
63
can be measured directly with a scanning electron microscope.1 where both membranes (of the same pore size rating) are viewed at the same magnification.

2). Polytetrafluoroethylene (PTFE) or other chemicallyresistant polymers cannot always be dissolved in an organic solvent. However. surfactants are added. The “capillary-pore“ membranes generally have porosities less than 5%.1 indicates that the “tortuous-pore” membranes are more porous-having a porosity over 75%.2:
Schematic of casting machine for MF membranes. and the membrane is dried.” Subsequently. Robert Gore developed a stretching process for making porous PTFE membranes.3 shows the nodes and interconnected fibrils of the porous membrane produced-GoretexTM . The belt passesthrough a series of environmental chambers usually containing water vapor at elevated temperatures. The more volatile solvents evaporate and the water vapor precipitates the polymer around the less volatile solvent which becomes the “pore-former. the fact that the latter are ‘/is the thickness of the “tortuous pore” membranes means that the flow rates are often comparable. the residual solvents are washed out of the pores. Typically.
CELLULOSIC WlrH
MATERIAL
SOLVENIS
Figure 2.64
Handbook of Industrial Membrane Technology
A look at the open area of the two membranes in Figure 2.2 is a simplified schematic of a casting machine which makes cellulose ester membranes.
. (not shown in Figure 2.
There are also a few cases where the “phase-inversion” process is accomplished by passing the belt through a liquid water bath to precipitate the polymer as is done with polyvinylidene fluoride.” Figure 2. Tortuous-Pore Membranes Phase-Inversion Process.’ Stretching Process.2 Gore’s process uses a paste-forming extrusion technique where the extrudate is stretched at a rate exceeding 2000% per second at an elevated temperature. Figure 2. a casting solution made up of the polymer and a multicomponent solvent system is metered onto a stainless steel belt or web. after the membrane is formed. The final length of the stretched product is over 50 times the original length. Most tortuous-pore membranes are made by a casting process known as “phase inversion.

The stretching creates elongated pores measuring 0.3: Photomicrograph
of PTFE membrane (GoreTexTM).02 by 0. to make CelegardTM. The low flow rates of this membrane limit applications primarily to battery separators and airvents.4:
Photomicrograph
of polypropylene
membrane
(CelgardTM). The resulting solution is then Cooled in a controlled way until the polymer precipitates around the solvent which serves as the "pore-former" at room temperature."
. an expanded polypropylene membrane (see Figure 2.4).
Thermal-Phase-Inversion Process. The process is called "thermal-phase-inversion. Still another approach to forming porous membranes out of Polymers not soluble at roOm temperature is to elevate the temperature until they dissolve in a selected solvent.2011 or 0.04 by 0.Microfiltration
65
Figure 2.4011.
A similar approach has been used by Celanese Corp.
Figure
2.

2 and 0.
.66
Handbook of Industrial
Membrane Technology
Tracor Hydronautics and Enka (Membrana) have used this process to make polypropylene MF membranes. Thermal neutrons in the reactor result in massive charged fission fragments bombarding the film.6) involving a nuclear reactor and an etch bath. Theoretically. Enka produces both tubes (5.
Figure
2.8 mm I.11) of polycarbonate or polyester is passed between two fission plates of U23S. heating to fuse and draw down to individual capillary diameters of 0.
capillary
membrane
(made
by
EN KA '5 thermal-phase-inversion
Capillary-Pore
Membranes
Many techniques have been proposed for making capillary-pore membranes including laser beams. photochemical etching. a fairly thin film (less than 20.3 Glass capillary arrays are now commercially available for laboratory use. The l'track-etch" process was first patented by Price and Walker . In the first step. electroforming. To date. The capillary-pore structure appears to result from controlled growth of a1umina crystals.D. a new alumina capillary-pore membrane has just been introduced by Anotec Separations Ltd./8.5 .5).4 Jl (see Figure 2.11.7 ). The bundle is then sliced to form thin discs with a regular capillary array./2.D. and ionotropy to orient anisotropic gel particles to form ionotropic-gel membranes. any dielectric material subjected to this treatment will be left with "damage tracks'l which will be more vulnerable to chemical attack than the bulk of the material.4 It consists of a two-step process (see Figure 2.6 mm O. It has been speculated that the solvents used were kerosene or various amines.5:
Photomicrograph
of
polypropylene process) .) and capillaries (1.II " tlons except the track-etch membrane produced by Nuclepore Corp. They are formed by assembling a large number of parallel glass capillary tubes. These particles leave "damage tracks" in the film where polymer chains have been ruptured and ionized (see Figure 2.D.) in pore sizes of 0.D. none of these methods has produced submicron capillary-pore membranes at a reasonable cost in the large areas suitable for industrial appl ica.6 mm O.5 mm I. Further.

Once “strikethrough” has been accomplished.6:
Track-etch process for capillary-pore membranes. The caustic will immediately “strike-through” the track and then begin to etch radially out from the track.1.7:
Damage-track in an organic P0lYrt-W. the higher the temperature and caustic concentration. the larger the pore size.
In the case of polycarbonate or polyester films.
. The longer the residence time in the etch tank.Microfiltration
67
Charged particles
A
&!if lxz
Pores
Figure 2.Lwhen etching a 12 /. approximately 10 /A of surface will be removed from the film. The final membrane thickness. the etch rate is the same in the pores as on the surface of the film. Thus. in Table 2.I pore. there is an upper limit to pore size of about 12 cc. the “damage tracks” are susceptible to attack by caustic (NaOH) solutions in the second step. Since films thicker than 15 to 20~ will not allow penetration of the fission particles. For example.A surface.
Figure 2. after removal of 12 /. when etching a 10~ pore. of is only 6 P. the starting thickness is 18 /.

In general. at a 0. there are 20 square feet of membrane containing 6 x 1012 pores. In an attempt to maximize flow rate and still maintain good strength. The probability of this occurrence is further reduced by collimating the fission fragments so that they bombard the membrane at angles between 0° to 29° with an average angle of attack of 10° (see Figure 2. there is a small probability that two tracks will be adjacent to each other forming doublet pores. However.6 Further. which retains the smallest bacteria. Obviously.9) . the pore density is increased for the smaller pore sizes in Table 2. (see Figure 2. quadruplets. can have a marked effect on the retention characteristics of microorganisms.Th is is of I it tie hel p when the pore size is large as in Figure 2. Therefore.1. triplets.Microfiltration
69
The density of pores (number of pores/cm2) is determined by the residence time and/or power in the nuclear reactor. at pore sizes of 0. pore size.1 JJ.2 = 50. adjacent and parallel to each other . are present.8).s. there is an upper limit of pore density above which the membrane becomes intolerably weak. the statistical probability of pore overlap on both faces of the membrane is very high. increase the flow rate by a facto tor of 2. There is an overwhelming probability that many doublets.and below. in a 10-inch pleated cartridge.2 JJ. this is the underlying reason why Nuclepore will never be able to make a "bubble-pointable'l cartridge out of this membrane.10.
Figure 2. etc.. the membrane thickness is reduced by 50% to 5 JJ. for each pore size.8:
Doublet and triplet
pores in capillary-pore
membrane.
. but not enough to form two distinct pores over a th ickness of only 10 JJ. pores diverge. Since the irradiation step is a random bombardment process. the L/D ratio for the pore is much higher 10/0. Even so. when one considers the number of pores in one cm2 (3 x 108). etc. Apart from damage to the membrane during manufacturing. Here. However. porosities above 10% result in low strength membranes. two 5 JJ.However. triplets. Higher pore densities result in higher flow rates. pore densities above 6 x 108 pores/cm2 require longer residence times in the reactor which tend to scorch the film with the longer exposure to radiation. two or three pores which run through the entire thickness of the membrane.

In fact. As shown in Figure 2. membranes were characterized by their ability to retain specific organisms 100%.
Figure 2.11. and this became the pore size rating for those membranes which retained Pseudomonas completely.r.70
Handbook of Industrial Membrane Technology
P~LYCARB~NATE FllM OPEN AREA: PORES/cm2: -15z 20-30 MILLION
rl-! I
i 00 100 I-
Figure 2. etc.9:
Angle of pores in capillary-pore membrane. Pseudomonas diminuta was thought to have a minimum diameter of 0. a significant variation in size has been documented: up to 30% in diameter and 67% in length.45 j. but tortuous pore membranes are more difficult.0 micron capillary-pore membrane. Since the first applications of MF membranes were in the filtration and culturing of microorganisms.
.22 cc. Likewise.
PORE SIZE DETERMINATION Challenge Tests It is easy to measure the pore size of capillary-pore membranes with a scanning electron microscope. Serratia marcescens was thought to have a diameter of 0.10:
Doublet pore in 5.

Other particles may also be used to test the membrane.0 pm
Serra tia marcescens
Pseudomonas-
Figure 2.
’ ‘. The maximum pore size may be determined by measuring the gas pressure required
. The number of spheres collected are counted with the use of a scanning electron microscope. Normally. Normally. automatic particle counters are used before and after the membrane. Monodisperse latex spheres (produced by Dow Chemical) in concentrations of 106-10’ particles per ml are sometimes used. Bubble-Point Test An easier test which is also nondestructive is the “bubble-point” test. some manufacturers are more conservative than othersstarving bacteria before the test to challenge the membrane with the smallest viabile organism possible.
Yeast
Coliform
1
5.
Every manufacturer still challenges tests measuring the “Beta ratio” (the number of organisms challenging the membrane f number of organisms passing through). a second membrane is used downstream of the test membrane to collect particles which pass.Microfiltration
71
.5 /J in size. a membrane is considered good if challenged with enough bacteria to plug the pores (typically lOs-10s organisms per cm2 of filter area) and none pass. However. This is much more tedious and less accurate than the bacteria challenge test where bacteria collected on the second membrane may be grown into colonies visible to the naked eye.11:
Challenge organisms for various pore sizes. Alternatively.
Red Blood Cell \. but the method is not as sensitive and is limited to particles over 0.

the membrane is completely wetted with liquid and then a gas pressure is applied to one side. The pressure at which this occurs is called the “bubble-point” pressure.Microfiltration
73
First.’ If the gas pressure is raised further.
PRESSUREc--PORE DIAMETER
Figure 2. As the gas pressure is gradually raised (see Figure 2. A pore size distribution may be calculated from these data: Indeed. (5) To calculate the mean-pore size. this is the “bubble point” and may be used to calculate the maximum pore size. (4) Continue to raise the pressure until the airflow through the wet membrane equals that through the dry membrane at the same pressure. ASTM method R316 constructs a line equal to one-half of the dry air flow through the membrane. The pressure at which the wet airflow intersects this line is the point at which the liquid has been expelled from one-half of the pores and may be used to calculate the mean pore size. an ASTM test-method’ outlines the procedure: (1) Measure the dry air flow and plot versus pressure (see Figure 2. more pores will begin to bubble until the liquid is displaced from the smallest pore and the gas-flow through the membrane equals that through a dry membrane.
mean. all liquid has been expelled from the membrane and the pressure may be used to calculate the minimum pore size. (3) Continue to raise the air pressure and measure airflow through the wet membrane as a function of pressure.
Procedure used in determining
maximum.13). At this point. from equation (I). (2) Wet the membrane and raise the air pressure until the first airbubble appears. The maximum pore size may then be calculated. the first gas bubble will emerge from the largest pore-where the capillary forces are lowest.12). and minimum
. there will be no gas flow through the pores until capillary forces are overcome releasing liquid from the pore. Obviously.13: pore size.

14 will be slightly low for tortuous pore membranes at pressures near the bubble point. Even for large areas. In practice. Yet. As previously noted. since the parameter Nrd2/Q in equation (4) is approximately the same for capillary and tortuous pore membranes.9 x IO-‘gm mols/atm/cm3 for N2 in water at 2O’C)
AP = the pressure differential across the membrane
Q = the tortuous membrane path length through the pore and across the
Figure 2. However. Membranes of equal porosity and thickness will show the same “diffusionalflow. the chart may be used for both. this increase is mitigated by the fact that the membrane support also offers additional resistance to diffusion. both capillary and tortuous pore membranes would be expected to show a diffusional-flow of about 0.74
Handbook of Industrial Membrane Technology
There is one complication in the bubble-point test referred to as “diffusionalflow. For example. diffuses across the liquid-filled pore in solution.64 x 1O-’ cm2/sec for N2 in water at 20°C) H = the solubility of the gas in the liquid (H = 6. it is easily distinguished from the much larger gas-flow at the bubble point. the diffusion path through liquid in the pore is reduced. resulting in higher diffusional-flow rates by equation (4). Equation (4) makes it clear
. the diffusional flow rate should only be 30 ml/min. most pore sizes of tortuous-pore membranes have approximately the same porosity and thickness. This is because the pore opening is generally larger than the pore neck (see Figure 2. Thus. and comes out of solution on the lowpressure side of the membrane. In practice. At higher pressures.15). the liquid will flow from the pore entry region until it reaches the smaller neck where it will stop unless the pressure exceeds the bubble point.” A small amount of gas-flow can result even through a pore is filled with liquid. The gas dissolves in the liquid in the pores at high pressure.” The chart was originally constructed for use with capillary pore membranes with a thickness of 10 cc. Therefore.” This is why the so-called “forward-flow” test advocated by some manufacturers is so misleading. (This can be readily measured with a liquid-filled inverted graduate or burette-see Figure 2. “diffusional-flow” is not even detected when small membrane areas are involved. They claim they can correlate bacteria retention with the diffusional-flow measured. Table 2.16).14 is a convenient chart to use in estimating “diffusional-flow.1 should be used to identify the density corresponding to the pore size in question. The estimates obtained from Figure 2. It is helpful to estimate the amount of flow expected due to diffusion: (4) J = m2 4 where (Dl-l) & L
J = gas flow rate per unit area of membrane N = number of pores per unit area d = pore diameter D = diffusivity of the gas in the liquid (D = 1.2 (u pore size has a density of 3 x lo* pores/cm2. For a 10 square foot pleated cartridge tested at 30 psi.1 ml/min/ft2/psi. a 0. Nnd2/Q is the key parameter in equation (4).

Lmembrane which will retain it quantitatively-if the membranes have the same porosity.
.
Valid for pressures below bubble point Assuming membrane thickness of IO )rrn
Pore size.Microfiltration
75
that a “forward-flow” test cannot distinguish between a 1.14: Diffusional-flow for various pore sizes and pore densities. and a 0. pm Figure 2.2 /J membrane which passesPseudomonas diminuta quantitatively.22 /.

Other Methods A number of other techniques have been used to determine pore size. The greatest ambiguity is that the volume of mercury penetrating deadended “pits” is registered along with that penetrating “pores. Because the contact angle (0) is 180°. but they are not accurate enough to be used as standard characterization tools. the pressure required to force mercury into the pores is almost seven times greater than the pressure required to expel water from the pores. Thus. The incremental volume of mercury added for each increase in pressure is measured precisely. Nevertheless. A rough estimate of an equivalent mean pore size may be made by measuring the porosity (E) and permeability (J) of the membrane.Microfiltration Mercury Intrusion Test
77
Mercury is a non-wetting fluid for most materials. The technique places as much membrane area as possible into a chamber which is evacuated.” Further. We speak of mercury “intrusion pressures”. Assuming laminar flow. the high pressures required may collapse some pores. Flow Permeability Test. the Hagen Poiseuille relationship may be modified for a porous membrane: 4 (5) J=_ 128~ L where J = volumetric liquid flow rate per unit area of membrane N= number of pores per unit area
d = pore diameter
AP = the pressure differential across the membrane cc = viscosity of the liquid P = the length of the pore
The porosity (e) of the membrane is the same as the void volume of the membrane and may be determined by measuring the difference between the wet and dry weights of the membrane. and pressure is required to force mercury into the pores-see equation (1). mercury porosimetry can be helpful in confirming the pore-size distribution obtained by other methods such as bubble-point. cos = -1. for a given pore size. Since
E d 2AP J= 32~1
. Mercury is then admitted into the chamber as the pressure is slowly elevated. these are quite high due to the high surface tension of mercury (476 dynes/cm).

312 shows that filtration time as well as challenge level is a factor. The classic method’ for determining the filtration efficiency of various air filters is to challenge the filter with a liquid aerosol of dioctyl phthalate (DOP).showing that MF membranes may pass microbes at high challenge levels (see Table 2. Table 2. may be significant.
. For example. Although the method has been used to characterize MF membranes for the filtration of gases. Wallhausser wrote several controversial papers. However. the use of this techniques is not widespread due to the tedious regimen required in gas adsorption measurements. II is related to the thickness of the membrane by a “tortuosity factor.
RETENTION
CHARACTERISTICS
Bacteria Retention and Bubble Point In 1974-1976.r cellulose ester membranes (water wet) have bubble-points between 45 and 60 psi.2). other data began to appear in the literature confirming Wallhausser’s findings.” The problem is that we cannot measure the “tortuosity factor” with any degree of certainty.” for example. Typically. The idea was that bacteria trapped in the neck of the pore subdivide during growth so that the new organisms emerge through the neck to the other side. A wealth of information about the size and shape of pores may be obtained from adsorption isotherms where the mols of nitrogen adsorbed on the membrane are measured as a function of pressure. Further. the theoretical water bubble-point of a membrane with a membrane with a maximum pore size of 0.“r” . inertial losses. Smoke DOP Test. Some researchers used the term “grow-through” to explain the results. BET Adsorption Data. which is only one-fourth of the theoretical value. Previously. Shortly thereafter. the length of the pore. calculation of the maximum pore-size from the bubble-point shows that there are some pores larger than challenging organism. the retention of a liquid aerosol does not correlate well with the retention of particles suspended in a liquid.2 I-( is calculated from equation (2) to be 200 psi. the aerosol particles are 0. Even the case of “capillary-pore” membranes where R is only slightly larger than the thickness of the membrane. the hysteresis effects make conclusions about pore-structure ambiguous. membrane manufacturers had insisted that their membranes were “absolute” and many were skeptical of Wallhauser’s data. especially “front-and-backface” losses.78 or
Handbook of Industrial Membrane Technology
(8)
d-
All the variables on the right side of equation (8) are experimentally measurable except for II. Eventually the manufacturers collected their own data which showed that no membrane is truly “absolute. Obviously.20 /. However.3 p in diameter. Most 0.

22 pm 0.25
manufacturers bubble
early
recognized They which
the
discrepancy
between into
theoretical into the account
and experimental tortuous
points. will attest.22 pm 0.2 pm
* PA .
membranes
surfactants
can dissolve
in the test
.
in shape.2 pm
Time. lower and
It is true that most It is conceivable the surface tension.2 Itm 0.
introduced
an experimental took
constant
(1) called a “shape-factor”
be justified chemistry from
supposedly rise equations.
The equation fact that 0.80
Handbook of Industrial Membrane Technology Table 2. other
an irregular
shape will yield a higher bubble the discrepancy tension fluid with wetting between
researchers treat their
explained with
experimental manufacturers these
bubble-points
surface
depression. equation
on surface adjusted
(9)
where p is the perimeter
p = wCos@
A of the pore and A is its area. for pores of equal point than a cirtheoretical and that
areas. hr 24 48 168 312 72 119 170 265 314 24 41 162 216 24 24 24 48 89 90
Bacteria Recovered O/L O/L 62/L 750/L O/L O/L 16/L 20/L 19/L I9 total I9 total 260 total IO00 total I total 0 1I total 1000 total 0 0
PA 5-248D
0.
pores are not cylindrical the capillary Indeed. agents. Thus. Note: All tests run to plugging (40+ psid).2 Ccm
CET-1 CED-1
0.22 pm
CET-2 CET-3 PA 538 PA 562 PA 595
0.” PA 5-2488
Rating 0. for noncircular
a shape factor of (1) is
cannot to:
as any good textbook pores.nylon 66. the one with cular pore! Still.3: Bacterial Penetration with Time12
Filter No. However. Input levels ranged from I O4to 1O5bacteria/ liter. CE = Cellulose ester.2 pm 0.22 pm 0.

From the above. For example.7 0.
of pore sizes in 0. it nevertheless demonstrated that a membrane can be made where a direct measurement of the pores (with S. the accumulation of bacteria on the membrane can reach these levels. where the membrane is in service for many months. the membrane has been “absolutely retentive.Microfiltration
81
However.8 /J in diameter. for all practical intents and purposes.
50
45 40 35 30 2s
IS 10 s 0.5 0.” There are few applications which subject the membrane to the challenge levels used by the manufacturers in their quality control (typically. it is impossible to postulate a depression in the surface tension of isopropyl alcohol from 22 to 5 dynes/cm. however. some will find a leak-path through the MF membrane.2 /.Ipore size membrane has at least one pore up to 0. it is clear that given enough time and bacteria. This is confirmed by mercury porosimetry data (see Figure 2.) agreed with that calculated from the bubble-point. a long history of successful membrane usage in the sterilization of fluids by filtration. Such a challenge provides more than one organism per square micron of filter area and often clogs the membrane.17).8 0. 108-10’ organisms per cm2 of filter area). the discrepancy between theoretical and experimental values exists for all test fluids. The lowest densities virtually eliminated all doublets and triplets and yielded bubble points equal to the theoretical value.9 I. It is apparent that the bubble point is telling us that a 0.2 p capillary pore membranes were made with extremely low pore densities. Though the low density rendered the membrane useless for applications requiring a reasonable flow rate. Indeed. There is. This explains why there can be more organisms in the filtrate than can be detected in the feed stream.E.
.0
Port-Size wn in
Figure 2.M.17: Distribution porosimetry).45 micron membrane (by mercury
To verify the theoretical equation for bubble point.. a series of 0.4 0.6 0. It is true that in ultrapure water loops.

it is recognized that larger microorganisms can deform.
.1813 demonstrates the transmission of normal deformable red cells through various capillary-pore membranes. “Tortuous-pore” membranes have 25 to 50 times more internal surface area for adsorption than “capillary-pore” membranes. they will not pass anything below a 9 /. a 5 p pore size will retain 60% of the 0. On the other hand. For example.005 ~1 Au colloids by “capillary-pore” and “tortuous-pore” membranes.18:
Transmission of red blood cells through capillary-pore membranesI
Retention by Adsorption Lukaszewicz et all4 has postulated an “adsorptive” mechanism to explain the excellent retention of bacteria by membranes with some pores larger than the microorganism. Most microbiologists are doubtful that Pseudomonas diminufa will deform to this degree. After hardening the red cells.0005 /J colloid. Figure 2. Indeed.I pore size. and the tortuous path also results in a greater likelihood of small particles contacting the pore-wail. under pressure.4 shows clearly that the “tortuous-pore membranes” retain particles much smaller than the rated pore size. “capillary-pore” membranes retain less than 1% of either. He contends that the adsorptive sequestrations of the membrane are more important than the geometric restraints of sieving. Davis et al” investigated the retention of 0. The author has seen data on the passage of yeast particles through pore sizes under 3 1-1 which has suggested passage by deformation.05 /J colloidal particles and 18% of the 0. bacteria will passthrough pores of smaller diameter due to deformation.05 and 0. Nevertheless. Table 2.
I
I
I
I
I
i
i
NORMAL IIC
HARDENED
I8C
1
I
I 4
I 6 FILTER PORE
8
I
DIAMETER
IO (p)
I
12
I
I4
I
Figure 2.82
Handbook of Industrial Membrane Technology
Retention of Deformable Particles Some researchers have argued that.

2 120:.3 6.5).Ovm Nuclepore 5.. For example. Since most particles are negatively charged...g..9 0.lum Nuclepore O...005 0.
.. Normally a crosslinking epoxy resin with cationic functional groups (e.
..3%%)
..7 46.7 A92
Colloid Size (urn) O... .
60 78
75
98 61
86 76
7
80
Polycarbonate.. BSA. in the data of Hahn et all6 (Table 2..5: Recovery of Poliovirus in Retentate or Filtrate after Filtration on Various Filters
SuspendinS r&am* Filter She @II) BSS
BSA(1%)
BE (O.lvm Cellulosic 0.45um Cellulosic l.
<5
<5
80
:
<5
7.05 9% 1.Ovm Nuclepore 1...
<5
<J
73
. BE.45 cellUlosc.45
Silver..... the retention of charged species may be enhanced without sacrificing the filtration rate by using charge modified membranes. product may be lost via adsorption on the membrane. Table 2.5 0...4: Retention of Au-Colloids on Capillary-Pore Membranes 0.2pm Cellulosic 3 ..2 ..) The recov-ery was improved from 5 to 76% by pretreating the membrane with a beef extract solution.. For example.. the “capillary-pore” configuration is best for fractionation of particles. For dilute process streams.. quarternary ammonium groups) is used to impart the charge.
l
.4Um Nuclepore 0. .. The “recovery” of this product may be improved by pretreating the membrane such that most of the adsorption sites are occupied..0ttmCellulosic
These results suggest that a “tortuous-pore” configuration is best for “cleanup” applications where the removal of all particles from the process stream is desired.2 8. polio virus adsorption on cellulosic “tortuouspore” membranes was significantly higher than that on polycarbonate “capillarypore” membranes.3
83
and Tortuous-Pore 0. . (i.Microfiltration Table 2.. Hanks balanced salt solution. On the other hand..e. The combination of electrokinetic adsorption with mechanical sieving is claimed to result in significant improvements in retention..
Silver. these membranes are usually positively charged.5
BSS. capillary-pore membranes have been used in fractionating silver colloids to improve resoltuion on photographic films. bovine albumin in distilled water..4 59. The virus recovery is low due to adsorption... 0.
Retention by “Charged” Membranes In certain aqueous streams.OOl?&) WI%1 BE
BE (0.3 46. 22 Cellulose. beef extract in BSS.

the membrane does not clog even with the processing of very high volumes. Table 2. Figure 2.20). For critical applications. For example.000 molecular weight cut-off UF membrane is required to remove pyrogens. (These membranes have an equivalent pore-size of 30 A or 0.‘s
. one cannot afford to guess when the adsorption capacity is used up. Their conventional 1. As expected.
Table 2. but the positively charged membrane shows better than 97% retention.6 ” compares the pyrogen removal efficiency for a positively charged nylon 0.2 p latex particles. with use. Further. the membrane begins to pass negatively charged particles.99% retention of Pseudomonas.2 p ULTIPOR Nh6TM membrane (without the charge) retained less than 50% of the Pseudomonas.003 or).2 ccfinal membrane filter by a factor of 2. and a conventional cellulosic 0.9% NaCl solutionat pH 6.84
Handbook of Industrial Membrane Technology
The most significant claim is for the removal of “pyrogens” (endotoxins). it is reported that a 3 p positively charged prefilter can extend the life of a 0.2 p pore size membrane. Just as a saturated ion-exchange resin is incapable of further retention of ions.‘s on their 1.6:
Pyrogen Removal by “Charged” Membranes
~1
Racedure: E.7 and passed through filters containedin a25mm filter holder.to 6-fold.22 p membrane. In both cases. However.2 p Ne6 PosidyneTM membrane show better than 99. if all of the particles retained are considerably smaller than the rated pore size (say less than 25%) and no larger particles are present. high flow rates (pressure drops) through the membrane may strip the negative particles away from the membrane (see Figure2.2 p pore size MF membranes will retain Pseudomonas diminuta. the conventional MF membrane shows no retention whatsoever. the filtration efficiency of these positivelycharged membranes begins to drop. As the membrane adsorbs more and more negatively-charged particles the adsorption sites (cationic functional groups) are used up. decreasing the load on the “absolute” filter and extending its life. a 10. challenges between 2 x lOlo and 7 x 1Ol2 bacteria per square foot of filter area were used. in some cases.
It is also claimed that positively charged 1. The small particles do not “bridge” and cause only a slight change in the pore diameter.2 p pore-size membrane. Data from Pall Corp. ZetaporTM. Similarly. coli purifiedcndotoxin wasaddedto 0.19The hydraulic shear forces have overcome the electrostatic attraction. This means that comparable removal of Pseudomonas can be achieved with more than 10 times the flow rate of a 0. a positively charged membrane of larger pore size than the final membrane may be used as a prefilter.19l’ shows the decline in retention for 0. The difficulty is that. Normally.

2 1: Zeta potential of positively charged and uncharged nylon membranes.”
.21). A new filter is then placed in the downstream position.
PH
Figure 2.86
Handbook of Industrial Membrane Technology
Alternatively.18
-25 \ -30 \. The three mechanisms of aerosol particle retention may be illustrated from the data of Spurny et aI” in Figures 2. The U-shaped curves are characteristic of the efficiency of aerosol particle collection as a function of particle size. The adsorption capacity of the upstream filter will be exhausted first. it is replaced by the downstream filter.5 (see Figure 2. Aerosol Retention The filtration of particles in a gas stream can be quite different from the filtration of the same particles in a liquid stream. “capillary-pore” membranes have a deeper minimum in the curves than do “tortuous-pore membranes. However. It should be noted that untreated nylon membranes have a negative zeta potential at pH values above 6.22 and 2.23. When the effluent from this filter begins to show passage. The fractionation wpability of “tortuous-pore membrane ” is enhanced considerably by using filters with a zeta potential of the same sign as the particles (usually negative). two positively charged membranes may be used in series. Negatively charged membranes are also available.

thereby increasing the number of particles captured (see Figure 2.23 show that 100% of all particles larger than the pore radii (R) are captured. Particles smaller than the rated pore size can still be captured by "inertial impaction" if they have enough mass to continue in a straight line when the flow streamlines bend to go through the pores.22) because there is less likelihood the particles will be able to follow the flow lines all the way through a narrow pore.24). The particle is simply too large to pass through a pore. The result is that the particles impinge on the rim of the pore or on the pore-wall near the poreentrance (see Figure 2. Very small particles will not be captured by interception or by inertial impaction since their mass is too small.
Figure 2. Inertial Impaction.23). Both Figure 2. Also. The smaller pore sizes increase capture by inertial impaction (see Figure 2. these
.88
Handbook of Industrial Membrane Technology
Direct Interception. higher gas velocities impart additional momentum or inertia to the particles. However.24: impaction. This mechanism is common to both liquid and gas filtration.22 and 2.
Latex particles captured on a capillary-pore
membrane by inertial
Diffusional Deposition.

smaller pore sizes increase the probability of particle impact with the pore wall because of the shorter path length (see Figure 2. In this case.25:
Silver
(Ag)
particles
captured
by
a capillary-pore
membrane
by
diffusional
deposition. For dry gases. electrostatic forces can also boost the capture of these small particles.221.Microfiltration
89
smaller particles have greater Brownian motion and greater diffusivities.251.231.
Figure
2. In gases. Lower gas velocities favor particle deposition because of a longer residence time in the pore (see Figure 2. the diffusivity is large enough to result in deposition on the pore-wall (see Figure 2.
.

Using this kind of filter media as a prefilter for the final membrane filter often provides an ideal combination (see Fig-
. Prefilters The use of positively-charged open-pore membranes as prefilters for a final “absolute” membrane has already been mentioned.)” Air sampling stations have used this membrane in the first stage. “tortuous-pore” membranes are preferred.86
Handbook of Industrial Membrane Technology
“U-shaped retention curves” are seldom seen in liquid filtration. It is well known that fibrous depth filters provide enormous dirt loading capacity compared with membrane filters. Membrane Dirt-Loading Capacity It should be obvious from the section on “Retention by Adsorption” that “tortuous-pore membranes”. The data were run at constant flow rate and show the increase in pressure drop across 47 mm filters (AP) with increasing volumes of fluid passing the filter. Further. with 25 to 50 times more internal surface area for adsorption than “capillary-pore” membranes should have higher “throughputs”.23 is much less pronounced. The particle diffusivities in the higher viscosity fluid are much smaller. the retention efficiency does not increase with smaller particle sizes. particles passing are collected on a tortuous-pore membrane in a second stage to simulate what is deposited in the lungs. “Tortuous(PTFE) membranes pore” cellulose ester (CE) and polytetrafluoroethylene showed the highest throughputs. “capillary-pore” membranes are preferred. their cost and “throughput” (volume processed before plugging) will often dictate the membrane of choice. On the other hand.by diffusional deposition.26. In general. For “tortuous-pore membranes” the minimum in the curves of Figures 2. a prefilter serves to take the load off the final filter. The best prefilters will have a high internal surface area (high dirt-loading capacity). One must use more area of the “capillary-pore” membranes to equal the same dirt-loading capacity of a “tortuous-pore” membrane. it has been found that an 8 /. As a result. That this is the case is shown in the throughput curves of Figure 2. Further. for some specialized analytical applications where fractionation of the aerosol particles is the objective. For example. the longer path length through the pore results in more and larger particles captured. if the objective is to capture as many particles as possible on the membrane.r “capillary-pore” membrane will collet air-pollution particles that are normally deposited in the upper respiratory tract (nasopharynx.22 and 2. those process schemes and operating variables which maximize filter life and throughput need to be considered to improve the economics of filtration. “Capillary-pore” polycarbonate (PC) membranes plugged most rapidly. Thus.
MEMBRANE
PLUGGING
AND THROUGHPUT
Once membranes with the appropriate retention characteristics have been identified. This is because the tortuous path results in more and smaller particles captured by inertial impaction.

Depth Pref ilter
~
Membrane Filter
Figure 2.
7ffROUGUPU7
-ML
Figure 2.Microfiltration
91
ure 2. Most of the contaminants are removed by the prefilter but the final membrane serves as the ultimate barrier trapping all particles leaking through the prefilter-including fibers which “sluff-off” the fibrous media due to “media-migration”.27: Depth prefilter over final membrane filter.26: Tap-water through-put for polycarbonate (PC) capillary-pore membranes and cellulose ester (CE).27).
. polytetrafluoroethylene (PTFE) totruouspore membranes.

The manufacturers give some guidance as to which prefilters should be used for a given membrane pore-size. if the prefilter shows a high pressure drop and the membrane a low pressure drop (Figure 2. Limiting
VOLUME Figure 2. Usually. but some users are fearful they may represent a similar health hazard. If the final membrane is plugged but the prefilter shows a low pressure drop (see Figure 2. Glass fibers are available in the finest diameters. The combinations providing the highest throughput are then evaluated with respect to cost of the media. The trend has been to use polypropylene or polyester fiber prefilters. Unfortunately. However. there is no standard rating system for prefilters. the test is carried out at constant flow rate and the rise in pressure drop is recorded versus the volume processed until the membrane plugs. the prefilter is doing a great job protect-
.
On the other hand. the prefilter is too coarse and a more retentive prefilter should be selected to protect the final filter. Melt blown or spun-bonded fibers are available in diameters near 1 p. The selection of the optimum prefilter/membrane combination is best done experimentally on the user’s process stream.28:
Batch Volume
Inadequate prefilter (too coarse) for final membrane filter.92
Handbook of Industrial Membrane Technology
The smaller the fiber diameter used in the prefilter. the greater the surface area for adsorption of particles and the better the retention of small particles. fine diameter glass and synthetic polymer fibers were substituted. asbestos fibers were recognized as the best prefilter media. The individual fibrils were smaller than 0.291. but unfortunately.281. when it was suspected that asbestos fibers presented a health hazard.01 p and they had a positive zeta potential. neither media equals the performance of asbestos. In the sixties. Multilayers of these media with appropriate calendering have resulted in surprisingly efficient prefilters. The search for the optimum prefilter media is often facilitated by testing the membrane and prefilter separately after the combination has reached the limiting pressure drop. The experimental procedure used by the author places various prefilters on top of the selected membrane in 47 mm holders which are run in parallel on the process stream to be filtered.

Microfiltration
93
ing the final filter.
.30).
Limiting*P --_-----__
AP
VOLUME Figure 2.
The optimum match between prefilter and membrane will usually show both prefilter and membrane plugged so that neither one has carried the complete load by itself (see Figure 2. This may be accomplished by adding additional prefilter area.30: Optimized prefilter/membrane
I Batch Volume combination. it is also possible that the prefilter may be too tight. but additional dirt-loading capacity is required.29:
Batch Volume
Inadequate prefilter area for final membrane filter. In some cases. a more open prefilter can be used before the tighter prefilter. Limiting AP -------_
AP
VOLUME Figure 2. If the final membrane has a large pore-size.

tests were run on serum at constant pressure showing the volume processed at various times. We call this “anisotropy”.r pore-size final membrane. an improvement in throughput is achieved. In Figure 2. Indeed. but usually at the expense of retention. More open membranes were used as prefilters for the final 0. The declining slope of the curves reflects the declining flow rate with time. The idea is to provide a built-in prefilter with the more open pore-size upstream of the finer pore-size. orientation can have a dramatic effect on throughput. PlOO is a glass-fiber prefilter normally matched with a 1 /.31.
. Obviously.45 /. When using discs of anisotropic membranes. One manufacturer puts a note in every box of membrane: “Use this side upstream”.r pore-size membrane.94
Handbook of Industrial Membrane Technology
A series of graded prefilters can provide the highest throughput. Serum through-put
46
60
&I
Jr
(MINUTES)
as a function of various combinations of pre-
Some MF membranes have a variation in pore size from upper to lower face (see Figure 2.32). the increases in throughput achieved by additional media must be weighed against media costs. At least one manufacturer intentionally makes an anistropic MF membrane.31: filters.
P
)-
I-
I
I
I
I
lo
TIME
Figure 2.

The resistance term consists of two parts: the resistance of the cake which accumulates on the upstream surface of the membrane ( Rc ) and the resistance contributed by the membrane itself (Rm ).33). If the limiting pressure drop is 30 psi (as in Figure 2. It should be obvious that two identical filters run at two different flow rates will show different pressure drops across the filter.P) across the membrane or filter (i. the filter run at 50% of the flow rate will process a larger volume before its pressure drop reaches the limit. In conventional filtration of particulates:
(10) J z ~p Rm + Rc The resistance to flow observed brane operating with of the membrane. the filtration rate per unit area or flux (J) will be proportional to the pressure drop (D.32:
Crossosection
photomicrograph-anisotropic
Effect of Filtration
Rate on Throughput
Irregardless of the type of filtration media. with ultrapure a specified water fluid ( Am) is easily determined by the resistance for a given mem-
and should
be a constant
and temperature. For the same volume processed.Microfiltration
95
Cross-section
of the BTS Polysulfone membrane.. For any type of pressure filtration. the driving force) divided by the resistance to flow.
. the filter run at 50% of the flow rate of the other will show 50% of the pressuredrop as well. and the same degree of pore-plugging. XSOO magnification tortuous-pore memo
Figure brane.e. it is universally true that lower filtration rates result in higher throughputs.
2. At half the filtration rate. the pressure drop will be one-half as well.

33:
Through-put curve run at three different filtration rates.6 for commercial slurries. R. If Rm is defined as above.)
AP = pressure drop s
. This is due to the continually increasing thickness of the cake and its compaction under the pressurized conditions of filtration. must also include the effect of pore plugging within the membrane. =
A
where ~1’ = constant dependent on properties of the cake w’ Vt = is the weight of dry particulates per unit volume of filtrate = volume of filtrate delivered or “throughput” = compressibility exponent of the cake. R. (s is zero for a perfectly noncompressible cake and unity for a perfectly compressible cake.1 to 0..96
Handbook of Industrial Membrane Technology
I
I
1 600
0
300
THROUGHPUT IN GALLONS
Figure 2. is more complicated. normal values range between 0. In the conventional filtration of particulates:
S
a'W"+AP)
p
(11)
R. it is a variable which increases as filtration proceeds resulting in a progressively lower filtration rate at constant pressure.
The resistance of the cake of accumulated particulates.

A +ap
If one assumes that the limiting resistance to flow is that due to accumulated particulates on the membrane or within the pores.WIicrofiltration P A = viscosity of the filtrate = area of the filtering surface
97
Combining equations (IO) and (I 1) AP (12) J= U’
W’Vt(AF)Sp + %l
A
Thus. Higher pressures simply increase the resistance to flow of the cake enough to offset increases in the flow rate due to the higher driving force (API. Particles smaller than the rated pore size may be captured on the membrane at high velocities. the flux declines as the throughput increases. equation (12) becomes:
(14)
J=
he!2
a’w’Vtu
It is interesting to note that for a “perfectly compressible” cake. the flux becomes independent of pressure.34 shows that equation (15) is a good approximation in at least one case. The variation of throughput with filtration rate has far-reaching implications on the most economical way to run conventional membrane filtration processes. whereas at low velocities. the throughput seems to increase more sharply for a decrease in flow intensity. For example.35). Figure 2. the total volume (in gallons) which may be processed before plugging (run to a set API is
. they may follow the flow streamlines more eqsily and pass through the membrane without capture. the throughput (total volume processed) per square foot of membrane area will be inversely proportional to the flow intensity or flux (in gal/min/ft2). Inverting equation (I 2)
(13)
-J-=
J
a'w'Vt (AP)~-'J. This may be due to an inertial impaction phenomena.
(15)
Equation (15) shows that as a first approximation. Rearranging equation (14) gives us the dependence of throughput on other variables. In other cases (see Figure 2. if a fixed volumetric flow rate (gal/min) must be filtered.I R. which equation (15) does not take into account.

increasing the fixed capital investment (in housings. gol/~mlnMt*)
to.
I
I
1.98
Handbook of industrial Membrane Technology
proportional to the square of the membrane area.) by two times (to increase the membrane area by a factor of 2) will increase the volume processed (to membrane exhaustion) by a factor of four.0 Normalized flow rote.34:
Tap-water through-put inversely proportional to filtration rate. Consequently. QoVmin
10.35:
Dramatic decrease in through-put with increasing flow intensity.
.
0. Two to three times the area of the “tortuous-pore” membranes can be pleated into a similar cartridge because the capillary-pore membranes are so much thinner. Incidentally. [J in equation (15) has units of gal/min/ft2] . this also explains how “capillary-pore” membrane cartridges can equal the throughput of “tortuous-pore” cartridges. etc.0
Figure 2.1
1. Thus. replacement costs of membranes will be reduced by a factor of 2.0 Flow Intensity.0
Figure 2.

with one stream in and two streams out. pleated cartridges of both “tortuous-pore” and “capillary-pore” membranes are now backwashed in several wineries. in some applications. Figure 2.1 to 10% with a stable flux would be impossible with TFF. “capillary-pore” membranes appear to be more ammenable to backwashing than “tortuous-pore” membranes. “Cross-flow filtration” (CFF) is compared with “through-flow filtration” (TFF) (sometimes called “dead-ended filtration”) in Figure 2. having the particles collected on the filter facilitates disposal and favors the more conventional TFF. Generally. where the product is the filtrate. As might be suspected. Nevertheless. like cell harvesting. Current pleated cartridge design gives the membrane good support in the forward-flow direction but poor support in the reverse direction.
. The sweeping action of the crossflow stream tangential to the membrane surface maintained a stable flux at constant concentration. However. Since the yeast particles are considerably larger than the 0. Figure 2. In plate and frame systems (see Figure 2. Particles removed from the process stream are not all deposited on the filter. Currently. A bleed stream (reject stream) may be removed to keep the concentration constant. The particles accumulate on the membrane and are disposed of with the membrane. Sometimes. most MF systems operate with the more conventional “throughflow filtration”-one stream in and one stream out. Though not recommended by the manufacturers. Of course. several users have extended the life of their cartridges significantly using low-pressure backwashing.40 shows the cross-flow concentration of yeast with a “capillarypore” membrane. internal pore fouling was nil. the objective is to extend the life of the membrane indefinitely.38. some process streams deposit particulates on the membrane that cannot be backwashed from either type. This technique has also been applied with moderate success to MF membranes. the disposal of the retentate stream becomes a problem.39). Cross-Flow Filtration Ultrafiltration and reverse osmosis have always used a fluid management technique known as “cross-flow filtration” to sweep away deposited particles from the membrane surface. In “cross-flow filtration” . A backwash fluid outlet is provided up-stream of the membrane to purge “backwash debris” from the system. larger membrane housings are used to serve as an accumulation reservoir for the filtrate. most are circulated as a retentate stream with ever-increasing concentration. pumping energy must be supplied to recirculate the cross-flow stream at velocities 10 to 100 times higher than the permeation rate (see Figure 2.36). the filtrate itself is often used as the backwash fluid.2 ~1pore size.Microfiltration Backwashing
99
The life of conventional filter media can often be extended by “backwashing”. In other applications. In some applications. The ability to concentrate yeast from 0. the product is the concentrated retentate stream.37 shows a relatively successful backwash experiment on “capillary-pore” membranes used to filter beer. Indeed. the membrane must be supported on both sides.

hrr. The flux declines rapidly at first.41: Cross-flow filtration of single cell protein suspension (0.83 WT.
I
I4
I
16
I
I8
I
20
I
22
0
I
24
2
TIME
Figure 2. %. CELL SOLIDS
DATA OBTAINED 0. vol % Figure 2. as boundary layer conditions are established. and then levels off with a diminishing rate of flux dewy.40: Cross-flow concentration of yeast with a capillary-pore membrane.
For suspensions of particles with sizes nearer to the pore size. some internal pore fouling will occur but at a greatly reduced rate.83 wt %) with 0.
0
RUN
24 Al 0.22 micron tortuous-pore membrane.
. Figure 2.41z2 shows crossflow filtration of a single cell protein suspension on a “tortuous-pore” membrane.Microfiltration
103
Yeast concentration.22 30 p YIL
YILLIPORE SlRAIQHt
FILTER PATH IO0 CHANNEL ml/mln
CIRCULATION RATE:
A P = 60 eaig
0
I
I
4
I
6
I
6
I
IO
I
12 .

any species gaining admittance passes completely through the membrane with no opportunity to block internal pores. faces the feed stream. we can examine what happens to the cake resistance (R. There is a finite thickness of cake that is never removed.
. The break in the line is due to the sweeping action of the flow tangential to the membrane surface./AP. with exceedingly small pores. the resistance of the membrane divided by the pressure drop. we would expect to see a straight line with a slope proportional to the concentration of particles in the feed stream.
Filtrate
Volume
V [ II -
Figure 2.104
Handbook of Industrial Membrane Technology
Referring to equation (13). The skin of the membrane. Unfortunately. As the filtrate volume begins to accumulate. there are no MF membranes currently on the market which have this degree of asymmetry (compare the cross section SEM photos of Figure 2. albeit at a slower rate.) in cross-flow filtration by plotting the reciprocal of the flux (l/J) versus the throughput volume (Vt). we obtain results like those of Figure 2.43. Since the smallest passage-way is at the pore entrance. the data shows a growing cake resistance. In cross-flow filtration. The only bottleneck is at the pore entrance which leads into an ever expanding chamber. The intercept with the ordinate (Vt = 0) is equal to R. but the ultimate solution is to develop MF membranes which have a high degree of anisotropy. UF membranes are so anisotropic that internal pore fouling is virtually eliminated. The flux decay due to internal pore fouling can often be relieved with backflushing or chemical cleaning as in Figure 2. and there is a growing resistance due to internal pore fouling. For example.42: Increase in cake resistance (reciprocal of flux) as a function of through-put for cross-flow filtration (CFF). In through-flow filtration at constant pressure.42.32 with those in Chapter 3).

In other words. the pore entrance resembles a long narrow passage more than an orifice.
2.43:
Restoration
of CFF flux
by backflushing
and chemical
cleaning.Microfiltration
105
Chemical Backflushing
Cleaning
Figure
2.
Figure brane. Plugging can occur by particles bridging across the passage.44) but the skin is too thick.44:
Cross-section
photomicrograph
of anisotropic
ceramic
MF mem-
.
Some inorganic membranes (alumina) have been recently introduced which are anisotropic (see Figure 2. some improvement in flux decay has been reported for these membranes operating with cross-flow on certain process streams. However.

106

Handbook of Industrial Membrane Technology

The development of a truly anisotropic MF membrane will be a major breakthrough for cross-flow microfiltration.

MEMBRANE

CONFIGURATION

The configuration of the membrane will obviously affect cost, ease of replacement, and efficiency of filtration. There are currently three primary configurations for MF membranes in industrial use: (1) plate and frame units (2) pleated cartridges modules

(3) tubular/hollow-fiber Plate and Frame Units

Figure 2.45 shows the internals of a stacked-plate membrane filter housing accommodating up to sixty 293 mm membranes with a maximum filtration area of 33 square feet (3.0 m2). Only “tortuous-pore” membrane discs may be used in such a stack. The reason is that the polycarbonate or polyester “capillary-pore” membranes currently available are very thin (typically 10 j.r thick) and easily pick up an electrostatic charge. It is almost impossible to load 293 mm discs of these membranes onto the plates. They cling to hands and wrinkles cannot be totally eliminated. It is possible to buy (in Japan) a heat sealed “sandwich” with polyester screens on both sides of the “capillary-pore” membrane (see Figure 2.46) which facilitates handling. The “sandwich” is sealed around the periphery to prevent lateral leakage. After loading such a stack of membranes, it is imperative that the unit be “bubble-pointed” to establish the integrity of the stack. It is very easy to damage a membrane in the process of loading, Even a pin-hole will show a bubblepoint of less than one psi. Further, every plate has gasket and “o-ring” seals which may be out of place. A small piece of dirt or grit in an “o-ring” groove can also create a leak which will show up as a very low bubble-point. Plate and frame units are also used for cross-flow filtration, Some units take sheet stock MF membranes while others work best with a preassembled membrane cassette-a sandwich of two outer layers of membrane sealed to an inner filtrate collection screen (see Figure 2.47). A cross-flow spacer is placed between the filter packets and stacked in a plate and frame arrangement (see Figure 2.48). In some systems, the cross-flow spacer is a screen, but the “flow-channel” spacer shown in Figure 2.47 is less prone to fouling. Pleated Cartridges Stacked plate units were the first successful configuration for large scale MF, in spite of three distinct problems: (1) (2) (3) Membrane damage during loading “O-ring” gasket seal problems High labor and downtime for membrane replacement.

It was early recognized that, if the replaceable element contained more membrane area (preferably in a protective package), the installation time could be greatly reduced. Unfortunately, the early “tortuous-pore” membranes could not be pleated without loss of integrity. Therefore, the first cartridges were a closed-end 2 inch diameter tube almost 2 feet long with one square foot of membrane wrapped around the outside and sealed longitudinally. This permitted the fabrication of a “bubble-pointable” tube. Unfortunately, the 22 inch tube contained only twice the area of a single 293 mm disc. For noncritical applications, pleated cartridges were fabricated showing less than full bubble point. For example, Figure 2.49 is a cartridge designed for filtration for deionized water. It contains five square feet of pleated 0.5 /A membrane prefilter and 0.2 @ membrane. Because the pleated membranes could not be “bubble-pointed”, a final core wrap of 0.8 /A membrane was incorporated to give the cartridge a bubble-point of around 1O-l 2 psi. The tighter 0.2 ccpore-size membrane could not be used for the core-wrap because of the high pressure differential it generated across the core wrap (only 0.2 square feet in filtration area).

110

Handbook

of Industrial

Membrane

Technology

Microfiltration

111

In the last decade, some manufacturers have been able to produce “bubblepointable” (45-50 psi) pleated cartridges. These include “tortuous-pore” nylon and polyvinylidene fluoride membranes. The 0.2 /A cartridges show quantitative retention of Pseudomonas diminura at challenge levels of 10” to 1 013 organisms. membranes, though pleatFor reasons enumerated earlier, “capillary-pore” able, will not yield satisfactory bubble-points and begin to pass Pseudomonas diminuta at challenge levels above 103-IO’ organisms even though the membrane area is 2 to 3 times greater. Most pleated cartridges are sealed inside housings with “o-rings”. There are external and internal “o-rings” designs, both of which are vastly superior to the old gasket seals which can easily unseat under pressure. Some pleated cartridges are encapsulated in a disposable plastic housing so that, apart from external hose connections, there are no seals to worry about. Pleated cartridges are considerably more expensive than the equivalent membrane area in 293 mm discs (2 to 4 times as much). Some users insist that the lower replacement labor does not off-set the additional expense. However, the trend is definitely towards pleated cartridges for through-flow filtration. The movement to pleated cartridges has been accelerated with the discovery that some exhausted cartridges can be reused after backwashing. Modified pleated cartridges have also been used in a cross-flow operating mode.” In the simplest modification, a sleeve is placed around the cartridge and the housing modified to withdraw the retentate stream from the bottom of the housing (see Figure 2.50). The cross-flow stream is forced into the pleats where it moves tangential to the membrane. Even with this simple modification, dramatic increases in throughput were observed in the dewatering of yeast (Figure 2.51) and activated carbon (Figure 2.52). The increase in throughput was found to be a strong function of the cross-flow velocity-as would be expected. The data of Figure 2.51 and 2.52 were run at a recirculation/permeate ratio of 16.

More recently, a line of cross-flow pleated cartridges has been introduced with a controlled spacing for the cross-flow path between the pleats (see Figure 2.53).% This alleviates the most serious problem with the simple modifications described in the preceding paragraph where bunching of pleats creates a nonuniform distribution of cross-flow channels.

The polyvinylidene difluoride (PVDF) membrane is anisotropic and might better be called an open UF membrane. In addition, the alumina membrane is somewhat asymmetric (see Figure 2.44). None of the other membranes are anisotropic.

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Handbook of Industrial Membrane Technology

All of the tubes and hollow-fibers are self-supporting except for the PVDF membrane which is cast on the inside of a porous polypropylene tube. In most cases, they are potted on the ends to separate the process stream from the permeate (see Figure 2.54). This bundle of tubes or capillaries becomes the replaceable module.

APPLICATIONS Many of the applications for MF derive from the excellent retention these membranes have for microorganisms. Indeed, the retention for bacteria or other organisms is often superior to what may be obtained from tighter UF and RO membranes. We may divide large-scale MF applications according to whether they utilize through-flow filtration (TFF) or cross-flow filtration (CFF). The former are more common. Through-Flow Filtration Applications Sterilization and Particle Removal (Pharmaceuticals). A great many of the drugs and solutions produced by the pharmaceutical industry or made up in the hospital pharmacy have to be both sterile and relatively free of particulate matter-especially if the product is to be injected into the bloodstream. For drugs and other products that will not withstand heat, sterilizing filtration is the only alternative. Tissue culture media, parenteral solutions, vaccines, human plasma fractions, antibiotics, diagnostic injectables are all sterilized by membrane filters. Even if MF has been applied by the manufacturer, the solution may need to be filtered again by the pharmacist due to contamination during mixing or reconstitution. A single microorganism can grow into thousands overnight. Serious and sometimes fatal infections can result. There is a growing clinical evidence that particulate matter in IV solutions can also be a serious health hazard. Direct blockage of blood vessels can occur. For example, the partial occlusion of retinal arteries can result in blind spots. Clot formation and emboli result because of the tendency of red cells to adhere to particles. Granulomas have resulted from inflammatory reactions where a particle is embedded in tissue. The U.S. Food and Drug Administration (FDA) has specified that:
Prior to filling, large volume parenteral drug products shall be filtered through systems having a final mean porosity of no more than 0.45 ~1.Process specifications shall indicate the maximum time during which a filtration system may be used. Such time lengths shall preclude microbial build-up to levels that may affect the biological quality of the large volume parenteral, and in no case shall it exceed 8 hours.‘s

Achieving low particulate levels in the final drug or parenteral solution usually requires filtration of the constituent water. The United States Pharmacopoea defines specifications and methods for production of water for injection. Ultrapure water systems in the pharmaceutical industry use reverse osmosis, ion exchange, and MF just as in the electronics industry (see below). Both industries seek to produce sterile/particle free water. However, in the pharmaceutical in-

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dustry, most water produced must also be pyrogen free. This requires charge modified MF membranes, or UF/RO. MF is also used to remove microorganisms and particulates from air and other gases used in the pharmaceutical industry. Some specific applications include: vent filters; filtration of compressed air used in sterilizers, filtration of air or nitrogen used for solution transfers or at filing lines, and filtration of air or nitrogen used in fermentors. Again, the need for “vent-filters” has been recognized by the FDA:
All stills and tanks holding liquid requiring microbial control shall have air vents with nonfiber-releasing sterilizable filters capable of preventing microbial contamination of the contents. Such filters shall be designed and installed so that they do not become wet. Filters shall be sterilized and installed aseptically. Tanks requiring air vents with filters include those holding water for manufacturing or final rinsing, water for cooling the drug product after sterilization, liquid components, and in-process solutions.25

It is clear enough why microbe retentive filters must be used as air-vents on fermentation vessels. The reason for the FDA requirement on storage or holding tanks is that the movement of liquids into and out of these vessels entails the flow of air or nitrogen to maintain a pressure balance; thus the vent provides an entryway for airborne contamination. For gas filtration, the membrane should be treated to make it hydrophobic if not already inherently so. Membranes made from polytetrafluoroethylene or polypropylene are already nonwetting, and wettable polymers are treated by the manufacturer to render them hydrophobic. If the gas filter is hydrophilic, water condensing on the filter or entrained by the gas will wet the pores and be retained by capillary action unless the differential pressure across the filter exceeds the “bubble-point” pressure. In this case, the filter is “blinded” by water and the flow is restricted considerably. In the case of air-vents, even if the filter has been sized adequately to handle the passage of air commensurate with liquid filing and withdrawal rates, the use of hydrophilic filters has caused the implosion or collapse of more than one storage tank, A hydrophobic filter passes air preferentially; indeed, a water intrusion pressure is required to force water into the pores. Of course, the opposite phenomenon occurs with hydrophilic liquid filters where air bubbles are present in the liquid process stream; it is called “air-binding”. Often a hydrophobic vent-filter will be used in conjunction with the main hydrophilic membrane filter to allow escape of accumulated air without permitting liquid leakage. Naturally, any filter used to sterilize liquids or gases must be sterile itself when installed. Again, quoting the FDA:
Solution filters shall be sterilized and installed aseptically. The integrity of solution filters shall be verified by an appropriate test, both prior to any large volume parenteral solution filtering operation, and at the conclusion of such operation before the filters are discarded. If the filter assembly fails the test at the conclusion of the filtering operation, all materials filtered through it during that filtering operation shall be rejected. Rejected materials may be refiltered using filters whose integrity has been verified.2s

Therefore, membrane filters which can be autoclaved or steam sterilized are

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in demand for these applications. If the membrane or cartridge cannot be autoclaved, other chemical sterilizing agents may be used (e.g., formaldehyde). Sterilization and Particle Removal (Beverages). MF has been used extensively in the filtration of wine and beer. It is also currently used in the clarification of cider and other juices. The objectives for beverage filtration are to obtain a biologically stable product with good clarity and no deposits. Since wine and beer are fermentation products, shelf life will be limited unless all yeast are removed. Unless beer is kept very cold, a few surviving pediococci, lactobacilli, or wild yeasts may grow and produce off-flavors in the beer. Though the juices are not fermentation products, they nevertheless provide the perfect media for bacteria growth. In all cases, it is important to produce asealed sterile product. There are three routes to biological stability: Heat stabilization Injection of chemical preservatives Sterile filtration Hear pasteurization is an effective means of biological stabilization, but flavor is affected. The bouquet of a good wine is highly susceptible to heat. The vintner goes to great lengths to protect his wine from oxidation, and heat accelerates it. In addition, elevated temperatures can result in precipitation of proteins to form haze. Pasteurization is still the most common practice in breweries in spite of some deterioration in flavor. A number of years ago, the brewing industry forecast a tremendous increase in demand for “draft-beer”, and membrane filtration was substituted for pasteurization. However, the “draft beer craze” never materialized and many breweries reverted back to pasteurization. Chemical stabilization has been accomplished by the addition of sorbates, benzoates, ascorbic acid, sulfur dioxide, or diethyl pyrocarbonate (DEPC). This method has not been universally accepted because of the public’s general aversion to chemical preservatives and undesirable side affects. For example, DEPC is very toxic until it breaks down into ethyl alcohol, carbon dioxide, and ethyl pyrocarbonate. The trend is toward the third approach of sterile filtration. Almost all wineries use it, though many breweries still prefer flash or “in-the-bottle pasteurization”. MF not only removes organisms but improves the clarity of the product. Juice producers are also beginning to explore membrane filtration. One large cranberry juice maker claims ceramic MF tubes pay for themselves in a year because of reduced labor, maintenance costs and waste. It is hoped that membrane filtration will also eliminate the need for heating the juice which changes its flavor. With the long-term trend of higher energy costs, it is expected that MF will eventually replace all pasteurization. MF has also made its debut in the Japanese soft drink industry for the filtration of liquid sugar. Figure 2.55 is a flow schematic for two parallel plate and frame backwash systems to remove yeast from the liquid sugar. The parallel lines permit backwashing without interruption of the process flow stream.

The development of the transistor in the early fifties and the first IC’s in 1959 made possible miniaturization on a scale never dreamed possible. The early solid-state devices were not as sophisticated as devices produced today.Microfiltration
119
Bottlers of carbonated soft-drinks are also experimenting with MF to remove bacteria and particles. Indeed. Deionized (01) Water. The accumulation of organic debris on the resins provides an almost ideal growth environment for bacteria which multiply at an exponential rate. these specifications are a trade-off between desired purity and practical purity and some water supplies will require more treatment than others to achieve these specifications. submicron particles can mask critical areas during etching and other key manufacturing steps. internal arcing. and a double pass to 0. can cause defects and reduce production yields.125 inch thick to which conductors were attached. and other unwanted changes in current carrying properties. For a typical feedwater with 200 ppm TDS (total dissolved solids). According to the commonly employed “Rule of Ten”. Conductor widths and spacing can be less than a micron. poor adhesion of photoresist coalings. Thus. undercutting. They quickly foul the resins and overrun the distribution system. This results in pinholing. Wafer-rinsing operations following wet-chemistry cleaning. there is constant mechanical breakdown of the resins
. including reverse osmosis (RO). If the water is not of sufficient quality. stripping and etching processes all require ultrahigh-purity process water as does process chemical-bath make up. It was clearly recognized that the presence of mobile ions such as sodium would produce leaky junctions and threshold voltage shifts in the final device. decreased wafer yields or subsequent device failure may result. Particle Removal (Semiconductor Process Fluids). The fluids that come in contact with photo masks and wafers during integrated circuit (IC) production are often contaminated with particles which. A single transistor consisted of a relatively large crystal of germanium or silicon 0. Thus.25 inch square and 0. the resistivity of the permeate from a double pass RO system is still insufficient (0.7% summarizes typical requirements for process water used in the manufacturing of high-density microelectronic devices. The latter is due to the removal of particulates which act as “nucleation sites” to cause premature carbonation loss. The heart of any ultrapure water system is ion exchange (IX).1 ~1. As shown in Table 2. is as effective as IX in removing ions to produce high resistivity water (IBM&m). it is not uncommon to have over a million components per circuit on a single wafer. Later. nearly every operation in IC manufacturing ends with a DI water rinse. No other process. Further. shorts. Obviously. advanced VLSI (very large scale integrated) microcircuits with one micron geometries are sensitive to particulate contamination greater than 0. DI water is crucial in the microelectronics manufacturing process. Greater consistency in taste and longer-lasting carbonation are cited as benefits. IX seems to aggravate the problem.1 to 10 M&m). all ultrapure water systems utilize IX usually in the form of separate cation and anion resin beds followed by mixed bed polishing columns.8. if not eliminated. The manufacturing processes were relatively insensitive to outside impurities.02-2 ppm. Today. Table 2. a single pass through an RO membrane can reduce the TDS to 2-20 ppm. it was discovered that 18 M&m water could still contain submicron particles and organics which introduce defects into the IC. because of the constant agitation and tumbling of the resin beads under forced-flow conditions.

Chloride.120
Handbook of Industrial Membrane Technology
themselves with the generation of resin fines and other particulates.
maximum
75 x 10-l mg/L two particles
Particle Living Total Total
count.01 to 0. As stated before.02 0.D.56.0 or more 10 or more 18.00
Water Purity Pure Very Pure Ultrapure Theoretical
Resistance at 25’C (Ma-cm) 0.
200 maximum x 10s3 mg/L mg/L mg/L mg/L 10 x lo-) 2 x 10-l 10 x lo-)
maximum Sodium.3
A typical ultrapure water system is shown in Figure 2.
Table 2. Thus.5 0.
level
Resistivity
(at 25OC)
of the time) minimum tppm) than larger
17 Ma-cm. Approximate Equivalent of Extraneous Electrolytes (ppm) 2 to 5 0.
maximum maximum organic solids.1 or more 1. the heart of the system is the IX demineralizers. maximum maximum
Table 2.
1 pm per milliliter one colony per 100 ml (ppm) (ppm) (ppm) (ppm)
carbon.2 to 0. organisms. there is a need for MF to remove all bacteria and other submicron particles after demineralization.
. All other components of the system are present to make the IX more efficient or to alleviate problems created by the IX columns-such as particle generation.S. (total).8:
Resistivity of DI Water as a Function of T.7:
Typical
Requirements for Process Water Used in Microelectronics Device Fabrication%
Parameter 18 Ma-cm (90% SiO.

_
.Microfiltration
121
E
LL
sl .

the combination of RO + IX is often more cost effective at TDS values below 300 ppm.S. the reaction of the acid with bicarbonate in the feedwater produces carbon dioxide which can further deplete the ion exchange capacity. RO operating costs are nearly independent of TDS in the normal range of operation.X. the water is typically placed in a storage tank which acts as a buffer during periods of peak usage.(ppm ae CaCOJ
700
I 10
Figure 2.
x
I
loo
200
300 4Oll 500 600 T. Therefore. and upgrading general system performance.D.
/
R. calcium sulfate. aluminum and magnesium. The foulants that most often plague cation resins are hydrous oxides of iron.D.) with and without reverse osmosis (RO). Anion resin fouling is generally caused by colloidal silica and high-molecularweight organic acids. manganese. grease.58) with a polishing mixed-bed deionizer is utilized by upgrade and maintain high quality 18 M&m water. Since high resistivity water is very corrosive and will leach ions from tanks and piping. one of the fastest ways to degrade this water is to allow it to stand still for any length of time. oil. most ultrapure water systems utilize RO to help prevent fouling of the IX columns (resulting in decreased exchange capacity). Forced-draft degasification removes the CO?. and suspended matter. RO can effectively remove many of these foulants inhibiting bacteria growth. Indeed some economic analysis show cross-over values as low as 100 ppm.0.57: Total operating costs as a function of T.
In any event. After passage through the cation and anion exchange columns (primary deionization).57. Ultraviolet steriliza-
. a continuously recirculating loop (see Figure 2. As shown in Figure 2.122
Handbook of Industrial Membrane Technology
RO with its attendant pretreatment (to prevent membrane fouling and hydrolysis) precedes IX to reduce the ionic load on these columns. copper. for ion exchange (1. Since acid addition is normally used to prevent scaling in the RO unit and to prevent hydrolysis of cellulose acetate RO membranes.S. Although IX regeneration costs are directly proportional to the TDS.

MF cartridges are still used for insurance at the point-of-use. particularly if cartridges must be replaced frequently. Chemicals. photoresist developers. When it is utilized only in the recirculating loop. When first installed on a 18 Mncm water line. Thus. The water filtered by MF is expensive. after drying. particle sloughing can be a problem with some cartridges.I. UF has also been used in the recirculating loop as well as at the point-of-use. which does not leach like polyvinyl chloride (PVC) piping.
MF is used to remove bacteria and other particulates both in the recirculating loop and also at the point-of-use.9% of the bacteria to keep the total population in the system within bounds. most support screens continue to slough submicron particles even after many gallons of throughput. In addition. particles can still build up on pipe walls and sluff off in time. a fair amount has already been invested in pretreatment by RO. Even though piping in the recirculating and distribution system is made of unpigmented polyvinylidene difluoride (PVDF). Even though rinsed in ultrapure water for long times. the rinse-down time can have a significant economic impact on water treatment costs as well as plant down-time. Even though the cartridges are prerinsed in ultrapure water by the manufacturer. the time for conductivity rinse-down or resistivity recovery of the filtrate can vary from as little as 2 minutes to as much as 60 minutes. initial exposure to ultrapure water results in further leaching of ions from the various components in the cartridge. Depending on the type of cartridge. water. Therefore. Figure 2.Microfiltration
123
tion is used to kill up to 99.
3rd POLlSlilNGLOOP Figure 2. Wet etchants and cleaners. all pleated cartridges degrade the water to a much lower resistivity.2 micron MF cartridges. etc. the MF membrane may do an excellent job of removing particles while the support screen is continually generating other plastic particles in the filtrate. IX. dopants and cleaning solvents are all a potential source of impurities
.49 shows the support screen which is downstream of the membrane. Selection of the optimum MF cartridge for this application must take into account conductivity rinse-down and particle sloughing as well as retention and throughput. This reduces the load on the 0.58: Flow schematic of recirculating polishing loop for D. dielectric polymers.

and piping. In these cases. Pb. However. Gases. a reject stream containing concentrated crud must be disposed of. In addition. U. they will not wet spontaneously with fluids having a surface tension greater than 32 dynes/cm. Since these membranes are hydrophobic.000 particles per cubic foot. Hg. Federal. valves. the units backwash every 3 hours at a back pressure twice that of the final forward pressure. For example. metal finishers. adhesion promoters such as HMDS) is particularly critical in high density devices. Both through-flow and cross-flow systems have been investigated. Cu. high throughputs are important. Several plate and frame MF units have been installed in Japanese nuclear power plants. aluminum and zinc. PTFE membranes or cartridges are ideally suited to aggressive chemical environments. magnetite.2 p before packaging. Ni. primarily iron oxides (hematite. etc.124
Handbook of Industrial Membrane Technology
which may reduce semiconductor device yields. state and local regulations place strict limits on the quantities of heavy metals which may be released to the environment. As. Sb and Zn. Particle Removal (Nuclear Power Industry).S. Previously. The units operate with automatic periodic backwash using accumulated filtrate. The controlled metals include: Ag. For this reason. Even high purity gasescan contain more than 1. particles can be generated from moving parts in compressors. One advantage of through-flow MF is that the RAD wastes are deposited on the filter which may be disposed of more conveniently. Typically. nuclear power plants have looked at MF and UF for removal of the crud. In this sense. and printed-circuit-board manufacturers are under mounting pressure to clean up their waste waters. The aqueous solution may then be used to displace the wetting fluid and the pores will remain water wet.) with small amount of copper.g. Cr.
. degradation of the resists can take place due to autopolymerization-generating “gel slugs”. the membrane must first be wetted with a miscible organic solvent such as isopropyl alcohol (27 dynes/cm). Cross-Flow Filtration Applications Removal of Heavy Metals. Cd. The backwashed crud is removed from the system as concentrated RAD wastes. Even though the photochemical manufacturers usually filter their products to 0. In addition.. the replacement of cartridge filters is a major undertaking because of the radioactivity hazard. the precoat material including the RAD wastes represents a sizeable volume of material which must be stored and treated for final disposal. insoluble particles generated from the photoresist pump at the point-of-use can recontaminate the material as it is being dispensed on the wafer. These radioactive materials. This includes aqueous solutions and even some organic solvents like ethylene glycol (44 dynes/cm). Therefore. There is a build up of corrosion products in the primary coolant loop of boiling water reactors (BWR’s) which must be removed. electroplaters. However. are referred to as “crud”. precoat filters have been used to remove the crud. Point-of-use membrane filtration and central process gas filtration are recommended to minimize the particle levels in the entire system. If there are air bubbles or entrained gases in the liquid. In either case. nickel. the need for controlling contamination in photoresist and related photolithographic chemicals (e. MF membranes or cartridges at the point-of-use ensure high production yields. Inert gases are used to dry silicon wafers. these back-wash systems are not dissimilar to cross-flow units using MF or UF-membranes. the membrane may need to be rewetted periodically.

2*r2”. the cross-flow membrane modules (usually tubular) concentrate the slurry to 2-10% solids.
of wastewater
CFF
to remove
precipitated
heavy
Industrial Laundry Wastewater. Further concentration to 20% solids is accomplished by gravity settling. the metal precipitate may be redissolved in concentrated acid to recover the metals in solution. Wastewater from all laundry sources accounts for approximately 10% of municipal sewer discharges.
. the levels of oil and grease. and if need be. Figure 2. In the United States. Although conventional methods of waste water treatment may use a similar pretreatment chemistry. polymers or electrolytes are introduced to flocculate and/or coagulate the colloidal suspended solids. However.59: Flow schematic metals by MF. industrial laundries alone account for over 10 billion gallons (38 million M3) of wastewater annually. Note the cleaning tank in Figure 2.5g2’ is a typical flow sheet for treating spent process water from printed circuit board manufacturing.59. The polymers and/or electrolytes neutralize the electrostatic surface charges which repel particles of like charge preventing coagulation. The laundry industry is a major generator of wastewater. Typically. the final solid/liquid separation by gravity settling is usually not as effective as membrane filtration. and/or catalytic reactions are utilized along with pH adjustment to provide the optimum precipitation.Microfiltration
125
M F may be used to remove these heavy metals provided pretreatment chemicals are added to precipitate the metals to particles of filterable size.
Figure 2. The chemical pretreatment step is crucial since it will affect the performance of the membrane and the resultant sludge volume as well as the contaminant removal efficiency. heavy metals and other organics exceed municipal discharge standards. the MF membranes are more prone to flux decay as discussed earlier. absorption/oxidation. MF membranes are preferred over UF because of the higher flux. Reduction/oxidation. In addition to high suspended solids and BOD loading. a filter press can produce a dry (30 to 50% solids) cake for disposal. Therefore. If the metals are of high value. To enhance settling. Usually. periodic cleaning or backflush is necessary to maintain the flux at a high level.

1 0.
Table 2. The supernatant plasma was then decanted and the red cells returned to the donorenabling plasma to be drawn from the same person as frequently as three times per week.100 1.
.
Plasmapheresis. safety and cost.5 3.l ND* ND ND 6-9
Nondetectable.1
<30 <lOO 40 none <O.9:
Cross-Flow MF of Industrial Laundry Water Raw
Assays BOD COD O&G Suspended Solids Pb Zn cu Cr Ni Fe Chloroform Benzene Perchlorethylene Toluene PH
l
Wastewater (mg/L) 1. Both chemical conditioning and the addition of adsorbing and/or absorbing agents are necessary to render the contaminants filterable. The effluent quality is variable and does not always meet discharge standards.1 0.1 5.5 9.1
0.0 1. The permeate may be recycled into the plant. In the past.126
Handbook of Industrial Membrane Technology
Until recently.l
<O. Most of this plasma is then processed to yield purified components such as albumin or anti-hemophilic factor (Factor VI II).1 0.7 0.000 4. the underflow is polished with sand or diatomaceous earth filtration.1 0.3 2.000 1.l <O.2
0.300 5. Pilot studies in 1982 have led to the installation of cross-flow MF in some industrial laundries. plasmapheresis was carried out with blood donors by collecting their whole blood in plastic bags which were then centrifuged to separate the red cells from the plasma.l
3.1 0. Table 2. The separation of plasma from whole blood by continuous membrane filtration represents an improvement over conventional centrifugation techniques in terms of efficiency.1 ‘6-9
<O.1 0.3
40
Recycling Memtek Criteria Effluent (mg/L) (mg/L) 30 100 10 none 0. Subsequently.9 0. Finally. the standard method for treatment of these wastes consisted of lime coagulation and flocculation with clarification by dissolved air flotation. the sludge is dewatered by vacuum filtration.l <O.l <O.1 0.g2’ shows the nature of the raw wastewater and the effluent from the membrane system.l <O.

Blatt and co-workers at Amicon developed a thin-channel cross-flow device for plasmapheresis.
50 ACD
Hematocrit Of 41t. We discovered that pore sizes above 0. Cross-flow filtration (CFF) is potentially a less expensive and simpler separation tool for continuous plasmapheresis.6 ~1 MF membrane at an acceptable flux.I preferred because of higher plasma fluxes. red cells and plasma could be readily separated with a 0.6 /.2 /J retained some of the higher molecular weight plasma proteins (notably albumin and lsG). Therefore. As shown in Figure 2. Stirred-cells were used to reduce the accumulation of blood cells on the membrane. Whole blood is continuously withdrawn from the donor and red cells are returned while plasma is continuously removed. these devices are expensive and quite complex. The time required for collection can be reduced considerably. However. anticoagulated at avsrag. since the donor’s red cells are not removed in bags for remote processing prior to reinfusion. The early attempts at membrane filtration of blood were disappointing. In the late sixties.Microfiltration
127
Continuous flow plasmapheresis is a superior alternative. However. 3o In this device.60: Plasmapheresis cross-flow velocity. requiring the attendance of skilled personnel during operation.
Continuous flow centrifuges are now available with disposable rotating bowls. but the stirring hemolyzed the red cells and the flux was low even at high stirring rates. Of 2. Large scale plasmapheresis of animal blood is also facilitated by CFF.
membrane
flux
and
hemolysis
as a function
of
.5 psig
CHANNEL VELOCITY IFPS) -IN 9 MIL CHANNELS
Figure 2. there is a limiting velocity above which the degree of hemolysis is unacceptable.4 and 0.6 cc were selected with 0. pore sizes between 0.8 ~1 occasionally leaked nonhemolyzed red cells while pore sizes below 0. the flux increases with the cross flow velocity. There are two major advantages: (1) (2) There is less chance of a mixup.60.

4 and 0. (For a more complete description of polarization effects on the membrane. the product must be isolated from the cells and other soluble components of the broth/growth medium. cytotoxic antitumor agents. The cellular components and the plasma are then processed separately to obtain the desired final products. In the case of extracellular products. hepatic coma. This avoids damage to the cellular components. rheumatoid arthritis has been treated successfully by cooling the plasma to precipitate and filter out these so-called “cryoprecipitates” before returning the plasma to the patient. and even fuels have been produced in carefully controlled fermentations. Plasma exchange therapy (PET). the offending substance may also be precipitated out of the serum at reduced temperatures (4%). Anion exchange resins have been used to remove bilirubin and bile acids. the compaction of red cells on the membrane at higher pressures results in lysis of the cells. In most instances. Plasmapheresis is currently used not only in the collection of plasma by blood banks but also therapeutically. numerous enzymes.6 /J pores. plugging by lipoproteins or protein aggregates is suggested. The most striking success of this modern form of “blood-letting” has been in the treatment of autoimmune diseases such as Good pasture’s syndrome and myasthenia gravis. For example. The products of the fermentation may be either extracellular or intracellular. see chapter 3 on Ultrafiltration. Sometimes it is preferable to purify the patient’s own plasma rather than to replace it. free fatty acids.) Membranes having smaller pore sizes could be operated at higher transmembrane pressures before hemolysis was observed-suggesting that the red cells are lysed through distortion into the pores. In addition. A key step in the fermentation cycle is the separation of the cells or cellular debris from the liquid phase of the fermentation broth-commonly called “cell harvesting”. some of the larger pharmaceutical houses are beginning to process pooled animal blood in a similar way but on a larger scale. it was discovered3’ that significant flux decay occurred not only with whole blood but with cell-free plasma. Interferon. Cell Harvesting/Washing. In the metabolic disease states. Apparently. and bile acids. novel antibiotics. the use or microorganism-based fermentations for the production of chemical products has greatly expanded. the offending substance is an antibody whose activity is directed against the patient’s own tissues. Though most of the applications for plasmapheresis are medically related. Since individual protein molecules should pass through 0. the abnormal solutes are generally of low molecular weight and/or protein bound. insulin. or the product of an antigen-antibody reaction. With intracellu-
. Charcoal sorbents have been used to remove aromatic amino acids. which accumulates to produce so-called “immune-complex disease”. In recent years. fine chemicals. The technique has also been used to detoxify patients from drug overdoses or poisons. work at Amicon for the Red Cross31 determined that for each crossflow velocity (fluid shear rate) there is a critical transmembrane pressure above which significant hemolysis occurs. removes harmful components in a patient’s plasma by returning his cellular components with either replacement or purified plasma. Clinical improvement has also been reported with patients suffering from disseminated breast cancer. and a host of other diseases and immunological/metabolic disorders.128
Handbook
of Industrial
Membrane Technology
Later. In other disease states. This is accomplished by treating the plasma filtrate with sorbents or further filtration.

Indeed. Ammonium sulfate is then added to the supernatant to precipitate the protein product from the media.61 shows data on cell viability as a function of time while recirculating through an open-channel plate and frame CFF system. plate and frame devices which utilize screen spacers between membranes are prone to accumulation of cells on the cross-members of the screen resulting in flow blockage. Since MF membranes are prone to internal pore fouling. CFF can usually recover more than 99% of the cells from the broth even when the cell density is the same as that of the broth.33 Mammalian cells. Figure 2. modules were only available with UF membranes.Microfiltration
129
far products. The generation of potentially harmful aerosols in centrifugation (when harvesting pathogenic organisms) is virtually eliminated with CFF due to operation in a closed system with vented tanks. the cellular biomass isseparated by centrifugation. thereby avoiding potential contamination of the product. Hollow fiber membrane modules have minimum hold-up volumes and can operate without blockage of the fibers if the I. In some cases. CFF offers the versatility of cell washing in a single continual process known as “diafiltration” (see chapter 3 on Ultrafiltration). even when processed at slow rates with relatively high g-forces. CFF appears to offer many advantages over conventional separation processes like centrifugation.47 are less prone to fouling. CFF is not plagued with losses by adsorption on rotary drum vacuum filters (up to 15%) or by phase changes in precipitation. new buffer or solvent may be added as filtrate is removed. There is some concern that the fluid shear forces present in CFF could damage fragile cells. This is followed by further centrifugation and dialysis to remove the residual ammonium sulfate from the protein product. Likewise. of the fiber is substantially larger than the maximum cell aggregate. However. Until recently. cell harvesting can be accomplished more efficiently with
. the cells must first be concentrated and then ruptured (lysed) to free the products. Batch sizes are unlimited with CFF whereas the rotor or bowl capacity of a centrifuge limits the volume which can be processed at one time. Further. but the resultant supernatants are often turbid. Viable cell recovery was over 95% for both mouse spleen and mammalian cells after concentrating 6 fold from an initial concentration of IO6 cells/ml. the cell debris must be separated from the soluble products. Subsequently. clear supernatants (filtrates) are obtained.d. vacuum filtration and precipitation/dialysis for this application. Continuous flow or batch centrifugation separates the cells according to their density difference. “Flow-channel spacers” like those shown in Figure 2. membrane cleanability with flux recovery is particularly important. are particularly fragile and shear sensitive. viable cell recovery has in general been higher than with centrifugation. hollow fiber. Of the various CFF modules available. Cell washing with centrifugation requires repetitive steps of pelleting and resuspension. Cross-flow filtration (CFF) can replace all of these steps with a significant improvement in recovery and yield. like those used in these experiments. Further. the addition and removal of precipitants is not necessary with CFF. Up to 25% of the cells are lost in the supernatant. The advent of polysulfone MF hollow fibers makes possible cleaning with acid and base and even autoclaving at 121” in some instances. Traditionally.

Theoretically. Indeed.
. concentration: Figure 2. UF may be used for the final separation of stroma from hemoglobin. the data imply that cells may be concentrated up to the close packed sphere density of 75 volume percent.5~10~ cells/ml.
9Q%Original Cell 85 Viability 60-
0
MouseSpleenCells
A Mammalian Cells
75 -
I IO
I 20
I 30
I 40
I So
I to
I 70
I a0
I 90
Time (minutes)
Initial cell density: OS. yeast.62 shows flux data for the recovery of hemoglobin from the stroma (lysed red cells) using a 0.100%. Figure 2.
flux: 80. if the red cells are washed free of all plasma proteins (using cross-flow MF) before lysing. in this case. presented earlier. Indeed.000 molecular weight cut-off (MWCO) UF membrane (pore size of 0.1.000. However. is an example of concentration of whole yeast cells. when the products of the fermentation are retained along with the cells by the UF membrane.40.61: system.08 /J) should be able to make this separation.2 /J pore size MF membrane. d-fold. Though the concentration of cells at the end of the run was only 10%. even a 300.000 daltons) from the plasma which form a retentive dynamic membrane on the surface of the UF membrane (see chapter 3-Ultrafiltration-on the formation of gel layers). However. a 1.000. pores below 0. 95.100 liters/m2/hr. M F is the only choice. the intracellular product is hemoglobin having a molecular weight of 64.000 MWCO UF membrane was able to pass the hemoglobin with no significant retention. Here.130
Handbook of Industrial Membrane Technology
asymmetric UF membranes which are not as susceptible to internal pore fouling as the symmetrical structure of MF membranes. However.1 /J showed an unacceptable retention for hemoglobin presumably because of trace amounts of gamma globulins (>160. co/i concentrations of 60% and 37% respectively have been harvested.
viable cell recovery:
Harvested cell viability during CFF in open-channel plate and frame
Figure 2. and E.

Recovery of hemoglobin from stroma using 0. including yeast-presses.63).. precoated vacuum filters and centrifuges. the tax on beer lost in this sediment is not refunded. However.FC/RCULAT/O#
RATE . This concept has been explored both for MF and UF membranes and will be discussed more thoroughly in chapter 3 on Ultrafiltration.Microfiltration
131
I
200 160 /60
UEMOGLO8A' RECOV&RY
%
K-RO# tC-/ P/n-40 0
S?TROMA (30 M/L PSI
(WUMAN o/AMV.. Suffice it to say that the continuous removal of growth-inhibiting metabolites from a cell culture along with the products of the fermentation can lead to significant increases in cell densities and product yield.//JO Poue = 18 ns
==/M/N
1 I
8 2 3
f 4
I. coagulated protein. The continuous separation of cells from the products of fermentation by membranes opens up the possibility of a “continuous membrane fermentor” (see Figure 2. so there is considerable economic incentive for recovery. Conventional processes.
\
I
I
5
6
789/O
20
30
I
Figure 2.2 micron MF in
The ability to concentrate yeast to high levels with CFF makes possible the recovery of beer from tank sediment (the so-called “gelgger”)..
. the use of tubular MF as a prefilter increased the UF flux by 100% on the average. MF is too expensive for use as a prefilter. there are some feed streams which severely foul UF membranes and where prefiltration with cross-flow MF is cost effective. Prefiltration for UF. fat. Continuous Cell Culture. In most cases.. are not able to recover beer of good quality from this residue. For example. In some cases. CFF is required because TFF (through-flow filtration) would plug the MF membrane immediately. casein fines. in the UF of milk or cheese whey. Tubular MF is recommended for this application since UF membranes retain some of the proteins and other flavor components in the beer along with the yeast. In some countries.GL>
RED (DPO
CELLS) 2 F/L ZER)
Rs. which is about 6 to 12% dry solids. In the concentration of milk.62: cross-flow. the products of the fermentation may be larger than the pores of a UF membrane necessitating the use of cross-flow MF. and microorganisms all cause severe membrane fouling.

34 The difference in vapor pressures provides the pressure driving force. If a temperature difference is maintained across the membrane. Scaling will be unlikely against a nonwettable surface. Since there is complete separation between the seawater and the distillate. For example.
Membrane Distillation. In evaporating solutions. seawater may be separated from its distillate by a PTFE (polytetrafluorethylene) or a PP (polypropylene) MF membrane. The liquids involved can be at any convenient pressure higher than the vapor pressure and less than the membrane water intrusion pressure. the only requirement is solution.63:
Flow schematic
of continuous
membrane
fermentor. and condensate separated and at suitable temperatures. if evaporationcondensation is reduced to the essentials only. The process is. with the vapor pressure of the seawater higher than that of the distillate. there is no entrainment and high purity distillate is produced.64). where they will condense (see Figure 2. However. The warm seawater is fed countercurrent to the distillate through a membrane module designed to permit only small temperature differences across the membrane (see Figure 2.34 By using a heat exchanger be-
. as in “flash” evaporation. heat consumption is reduced by using an increased number of effects or stages.132
Handbook
of industrial
Membrane
Technology
STERILE MEDIA FEED
STERILE
CONCENTRATED
FILTRATE
Figure 2.65). each of which must normally be at a different pressure from the others. a multieffect evaporation process with liquid flowing from pore to pore (or “stage to stage”) but with the following advantages. evaporation will occur at the seawater side of the membrane and vapors will flow through the pores to the cooler surface. vapor. These requirements can be met by a porous hydrophobic membrane which excludes liquid from the pores but not vapors. in effect.

Feed Distillate Concentrate
Figure 2. This recovered heat can be 80% or more if low temperature differences are maintained and lower fluxes are acceptable.
hydrophobic microporousmembrane
Figure 2.64: Trans membrane distillation with hydrophobic MF membrane. part of the evaporation heat can be recovered.Microfiltration
133
tween the distillate and the seawater.
.65: Flow schematic of trans membrane distillation (TMD) with heat exchanger (HX) for recovery of heat of evaporation.

It is no coincidence that the United States is a world leader. He discovered that cellulose acetate (CA) is semipermeable to seawater electrolytes. but also in those technologies that are direct spin-offs-namely UF and gas separations. paved the way for the first anisotropic UF membrane in 1963. but both had impractical filtration rates (flux). Charles E. The breakthrough. The hydraulic permeability was low and the pores were easily plugged. The first work in RO toward water-desalting was undertaken by Prof. Before 1960.* The diffi136
. not only in RO. which resulted in an anisotropic RO membrane in 1959. Department of the Interior set up the Office of Saline Water (OSW) and committed substantial financial resources to the development of various separation processes for water desalination. The result was one of the most promising large-scale processes for inexpensive desalination of seawater and brackish water-reverse osmosis (RO). A significant portion of these funds was dedicated to the development of membranes for desalination-with funded programs continuing for over two decades (1950-1973). Porter
INTRODUCTION The beginnings of ultrafiltration (UF) are coincident with that of reverse osmosis (RO) around 1960. Despite the fact that the term “ultrafiltration” first appeared in the colloid literature toward the end of the last century and Bechhold. the evolutionary development of both occurred in parallel. In the years immediately following the end of World War I I.S. Though UF membranes are porous and RO membranes are not. membranes showing the retention properties of RO and of UF were available. in 1906 produced collodion membranes with pore sizes below 0. The U.3 Ultrafiltration
Mark C. Reid’ at the University of Florida in the mid-1950’s. the United States Government became concerned about shortages in water before the end of the century.01 micron. these membranes were little more than laboratory curiosities.

10. but the water flux decreased to unacceptable levels. The significant breakthrough was that the annealed membrane water-flux was 200 times greater than Sourirajan’s CA films and 5 times greater than the annealed S & S membrane. but found that below 6 /. Prof. Thus. Eventually the Mg(ClO& diffused out of the pores into the water bath. the value was 666 times greater than Sourirajan’s CA films. Sidney Loeb.) with astonishingly high filtration rates. Los Angeles (UCLA) obtained the same results-94% salt rejectionbut with water fluxes even lower than those reported by Reid.7.66. Suddenly. Dobry3 in a literature search. When the product of the water flux and the total membrane thickness was calculated. Loeb postulated the existence of a dense skin less than 1 /. Loeb heated the membranes under water (annealing) to temperatures between 80” to 9O”C. was to partially replace the aqueous Mg(ClO. the films had defects which passed salt and were too fragile to handle. In the early 1960’s. The CA precipitated around the Mg(CIO& leaving a porous film of CA. of the Massachusetts Institute of Technology and the founder of Amicon Corporation. thereby increasing the salt rejection from 0 to 92%. cast UF membranes from polyelectrolyte complex hydrogels in 1963. but the membranes were too porous and had no rejection for salt. The solution to the problem. a UCLA graduate student. the techniques utilized in the fabrication of asymmetric RO membranes were discovered to be applicable to the production of high-flux UF membranes with pores in the range of IO to 1000 a (see Figure P.l in the Preface). Michaels.& solution with acetone. The reason for this breakthrough resided in the asymmetric structure of the membrane. The most obvious explanation was that the effective membrane thickness was much less than the total membrane thickness. Sourirajan’s partner. This casting dope was spread as a thin film on a glass plate which was then plunged under water. suggested by Lloyd Graham. Loeb uncovered the work of a French investigator.2. He attempted to increase the ratio of CA to Mg(ClO& but the casting solution viscosity was too high.Ultrafiltration
137
culty was that the water fluxes were too low to be interesting. water.1 . UF on an industrial scale became practical. Dobry dissolved acetylated CA in an aqueous solution of magnesium perchlorate.4 When immersed in ice water. the substrate provided mechanical strength and the thin skin minimized the resistance to hydraulic permeability through the membrane.I in thickness supported by a relatively porous substrate. Reid and coworkers attempted to cast thinner films. Loeb then annealed these membranes to 80°C to yield a salt rejection of 99%. a solvent for CA.I in thickness. Her work provided the clue which Loeb was looking for. acetone. Similtaneously. A joint development program between Amicon and
. Loeb repeated Dobry’s recipe. Macromolecular separation could be carried out at modest pressures (less than 6 atm. Loeb’s standard casting solution contained CA. Mg(ClO&.0 and 1. and Mg(ClO& in the weight percentages of 22. and unbeknownst to Reid. a high flux membrane was obtained albeit with low salt rejection (typically 5% or less). began to experiment with laboratory UF membranes (not fully asymmetric) made by Schleicher and Schuell of cellulose acetate (CA). it became possible to remove salt from water (95 to 98%) at pressures of 50 to 75 atmospheres with flux values of 10 to 15 gallons of product water per day per square foot of membrane area (GSFD). For the first time in history. Alan S. Sourirajan at the University of California.

.1 to 1 . This thin skin permits high hydraulic permeability while the more open/porous substructure (typically 125 .u in thickness) provides good mechanical support. The solute molecule sees an ever-widening pore channel with no restrictions or bottlenecks leading to entrapment. In addition.1 is a cross section of a typical asymmetric (anisotropic) UF membrane. Tubes. it easily passes through to the other side of the membrane. once a solute molecule gains entrance into the pore.000.1:
Cross-section photomicrograph
of asymmetric
UF membrane. plate and frame units. By 1965.
Figure 3.138
Handbook
of Industrial
Membrane
Technology
Dorr-Oliver began to search for other polymers suitable for casting asymmetric UF membranes.u in thickness.
MEMBRANE
STRUCTURE
AND FABRICATION
Figure 3. The prominent feature of these membranes is the thin skin on the surface-usually 0. Hollow fibers were also developed during this decade and a whole host of module configurations. the pore configuration virtually eliminates internal pore-fouling. and spiral-wound modules became available. The ten-year period between 1965 and 1975 was a period of intense development where chemically and thermally resistant membranes were made from polymers like polysulfone (PS) and even polyvinylidene difluoride (PVDF) in molecular weight cut-offs (MWCO) from 500 to 1.000. the first laboratory-scale UF membranes and cells appeared on the market. Since the minimum pore size is at the membrane surface.

there are inorganic UF membranes made from zirconium and aluminum oxides.
Cross-section of UF membrane
cast on spun-bonded
polyethylene
As mentioned earlier. polyarylsulfone (PAS). aliphatic. Cellulose acetate and polyelectrolytes were among the first synthetic polymers to be used for UF membranes. the line of demarcation between UF and RO usually refers to the tightest U F membranes as those able to pass salts but retain the simple sugars. and polyvinylidene difluoride (PVDF). polyacrylonitrile (PAN).
Figure 3. Today.2: backing.
. In addition.Ultrafiltration
139
Additional strength is sometimes provided by casting the membrane on a spun-bonded polyethylene or polypropylene backing (see Figure 3. In fact. polycarbonate (PC). polysulfone (PS). the procedure used by Loeb and Sourirajan to make RO membranes from CA produces a very tight UF membrane (only 5% NaCI rejection) prior to annealing. UF membranes are made from a wide variety of chemically and thermally stable synthetic polymers including polyvinyl chloride (PVC). polyimides (PI ). and aromatic polyamides (PA).2).

46 is a schematic of a “drum casting machine” which carries out steps 2-4 in the above procedure. In the latter case.. or plasticizers (for drying) as needed. water vapor) from the vapor phase to the polymer solution. With increasing supersaturation.4 should be closed. at the outer surface of the casting solution. Once the skin is formed. (6) Drying (optional) at appropriate temperature/humidity. all of the above are solution cast or spun from a polymer solution. Strathmann describes the difference between a vapor-phase Precipitation process (Figure 1. The solution must then stand until all air bubbles are eliminated.140
Handbook of Industrial Membrane Technology
With the exception of the inorganic membranes. Procedure for Casting Flat-Sheet Membranes In general. (2) Metering the casting solution onto a casting belt.22) as used in MF membrane formation and the liquid-phase precipitation process (Figure 1.3). (4) Leaching remaining solvent out of the film. Since this is a relatively slow process. the procedure for casting anisotropic UF membranes can be thought of as a six-step process: (1) Preparation of the casting solution [dissolving the polymer in an appropriate solvent(s) with or without various additives] . PAS and PVDF. (5) Applying surfactants. For the more chemically resistant polymers like PS. Consequently. Thus. it becomes a barrier to further diffusion of water into the bulk of the casting solution. If there is appreciable solvent-loss with time. the critical nucleus size decreases and the individual crystals are smaller (more finely dispersed). In an earlier paper. In Chapter 1. Step number 3 in the above sequence is responsible for the formation of an anisotropic membrane. described here. solvents like dimethylformamide (DMF).23) used in UF membrane formation. in contact with liquid water. the degree of supersaturation is extremely high and the density of nuclei is high resulting in a finely dispersed structure which corresponds to the final membrane skin. wetting agents.g. the casting solution can be made up the night before the run and put on a roll mill in sealed glass jars which are heated by infra-red lamps. Likewise here. a rotating drum. or a reinforcing web/fabric. the degree of supersaturation is significantly lowered and the precipitate becomes increasingly coarse. precipitation of polymer is also slow resulting in fairly large Pores in the membrane. Figure 3. the rate limiting step is the slow diffusion of Precipitant (e. or dimethylacetamide (DMAc) are required. the pore size increases as we move from the skin into the substructure (see Figure 3. (3) immersing the freshly cast film into a nonsolvent liquid bath (normally water) until all of the polymer has precipitated/solidified and most of the solvent has left the film. For small batches.
.’ Strathmann uses the analogy of crystallization from supersaturated solutions. bringing liquid water in contact with the polymer solution results in catastrophic precipitation under supersaturated conditions. dimethylsulfoxide (DMSO). the solution trough shown in Figure 3. the rate of precipitation is higher. In the former case.

too much penetration will result in a considerable resistance to flow in the backing material. it must be borne in mind that high resistivity water will result in faster leach rates of solvent from the gelling casting solution which may affect the pore size of the final membrane. to
. On the other hand. With a UF caster. resulting in membrane delamination. This cost difference is accentuated by the fact that an MF caster must often run for an hour before any product comes off the machine. DI water with an 18Ma resistivity is the best standard. A 10 foot long casting machine with gel tank for casting UF membranes may run as fast as ten feet per minute-an order of magnitude faster than the MF caster. if they are to be dried. A 60 foot long casting machine with environmental chambers for casting MF membranes may run as slowly as one foot per minute. Therefore. the bond between the membrane and the backing will be weak.4 shows the rinse tank adjacent to the gel tank with continuous rinsing of the cast/precipitated product. Otherwise. Though Figure 3. If there is too little penetration into the fabric.142
Handbook of Industrial Membrane Technology
The doctor blade on the solution trough is adjusted so as to meter a coating thickness on the drum or fabric between 100 and 500 ~1in thickness. UF casting machines have a much higher rate of output.000 MWCO UF membrane. If bacteria slime or turbid water is noted. Further. The author has successfully cast many large batches of UF membranes using tap water in the quench bath. the tank should be drained and cleaned. However. the speed of the casting process is determined by the gelation time of the membrane and the physical dimensions of the gel tank. Because the gelation time for casting UF membranes with a liquid precipitant is much less than that for MF membranes with a gaseous precipitant. UF membranes are often less expensive to make than MF membranes. Plasticizers are necessary for some membranes. Indeed. in practice it is a separate operation. It will be noted that the cast film is immersed in the nonsolvent (water) of the gel tank immediately and no roller touches the surface of the film until it is completely gelled. continuously to minimize the buildup of solvents or other contaminants in the quench bath. product is produced almost immediately and feedback is rapid. Depyrogenization may be accomplished by filtering water from the ion exchange columns with a 10. a greater thickness is required. A separate tank may also be used to apply wetting agents or plasticizers. If the viscosity of the casting solution is such that it rapidly penetrates into the fabric. the process is much simpler than that required to make MF membranes. Considerable attention must be given to obtaining the proper match between a reinforcing fabric and the viscosity of the casting solution. contrary to popular myth. which is often thicker than the membrane itself. the use of depyrogenated water is required. When producing membranes for use in the pharmaceutical industry. However. The feedback for adjustments to the process is greatly delayed resulting in considerable waste material before conditions in the environmental chambers are set. It is advisable to use 01 water in the gel tank. this often results in membrane product which is discolored (iron oxide) and variable with the seasons as the tap water quality changes. This is necessary because of the long residence times required in the rinse tank to remove final traces of solvents from the membrane. pyrogens will be incorporated in the cast membrane and some will leach out in the filtratewith use. Water in the gel bath must be replaced periodically or better yet.

1POLYMER YIELD
STRESS
*p= ( -+)
SUSTAINABLE
i
MEAN PORE SiZGyr-
DISTANCE FROM SKIN SURFACE
Figure 3. Figure 3. Hydrophobic polymers like PVDF can be dried without irreversible damage to the pores. The pores of UF and RO membranes are small enough that the capillary stresses at the air-water interface can exceed the yield strength of “soft” polymers.
Y=
the surface tension of the liquid in the pores (‘y= 72 dynes/cm for water). a 15% solution of glycerin or polyethylene glycol in water or isopropanol can be used to saturate the pores prior to drying. “Hard” polymers with the larger pore sizes can also be dried without ill-effect. but must be rewetted to initiate flow.
(I)
where AP = the pressure difference across the curved interface between the liquid and gas.
. This not only reduces the surface tension. The “soft” polymers can be dried if the surface tension of the liquid filling the pores is reduced.
r = radius of curvature of the interface (equal to the radius of the pore for perfectly wetting or perfectly nonwetting pores). but the glycol remains behind after drying-serving as a plasticizer and wetting agent in the dry membrane.5 shows the capillary stress in the small pores of a UF membrane during drying. This may be accomplished by replacing the water with a low boiling organic like alcohol or acetone which has a low surface tension and can be evaporated without creating large capillary stresses.5:
Capillary consolidation stress in a UF membrane. Alternatively.Ultrafiltration
143
prevent collapse of the pores during drying. This latter method is used by most manufacturers who ship “dry“ membranes.

Porous polyethylene. All other steps are similar to those taken with sheet stock membranes.
(A)
0)
Inner Diameter Casting Tube
Membrane Tube
Immersion Tank Ice Water (O-4°C)MEMBRANE TUBE CASTING
Figure 3.
Casting Variables for Tubes and Sheets In Chapter 1. Strathmann used a ternary phase diagram (Figure 1.144
Handbook of Industrial Membrane Technology
Procedure for Casting Tubes The same six steps in the procedure for casting sheet stock can be applied to the casting of membrane on the inside of a porous tube.6:
Schematic of tube casting operation. The challenge is to find an inexpensive porous tube which has sufficient strength to withstand 100 psi pressure internally for long periods of time. Dr. As the tube is lowered further. Figure 3. In this case.13) to describe what happens when a nonsolvent like water is added to a homogeneous
. it is immersed in a water bath which precipitates the membrane. polypropylene.6’ is a schematic of a tube casting operation. and fiberglass-reinforced epoxy support tubes have been utilized. the tube is lowered over a casting bob which meters the casting solution uniformly on the inside wall of the tube.

if we assume that a membrane having a larger porosity overall will have a higher porosity in the skin as well as in the substructure. Path AD leads to a tight RO membrane because solvent leaves the film faster than water enters.. Path AE produces an open UF membrane because water enters the film faster than the solvent leaves. there is less volume of solvent present resulting in a lower porosity and smaller pores. and that the higher porosity is associated with higher permeability and larger pores.Ultrafiltration
145
solution of polymer. nor does it distinguish between the dense skin and the porous sublayer. The tool is limited because it does not provide direct information about pore sizes. the model can be very instructive.7: Ternary phase diagram for formation of UF membranes from cellulose acetate.’ The overall porosity of an anisotropic membrane is determined by the polymer content of the casting solution and by the relative rates at which nonsolvent enters and nonsolvent leaves the casting solution (i. the simplest way to adjust the overall porosity is to adjust the polymer content of the casting solution unless the resultant viscosity is out of
. the precipitation path in the phase diagram). However. and CA precipitates around a larger volume of solvent which acts as the “pore-former”.7’ is a ternary phase diagram for a castmg dope made of CA and acetone. it shows several different paths of precipitation for porous UF membranes and dense RO membranes. This is a very useful tool in understanding how we can alter the casting variables to form membranes of varying pore size and permeability. Obviously.e. When CA precipitates. Figure 3.
CELLULOSE ACETATE
rate than
solve
phase SOLVENT WATER
Figure 3.

this technique has been used to form MF membranes from polyvinylidene difluoride (PVDF). a lower porosity and tighter membrane will result. Water was one of the earliest additives incorporated into cellulose acetate casting solutions resulting in more open membranes. ZnC12. In addition to polymer content additives. Eirich et aI9 discovered that regardless of the nature of the polymer. without incipient precipitation of polymer. and polyacrylonitrile/poIyvinyl
.8’ shows the effect of additives in the casting solution (formamide. Loeb used ice water (see Figure 3. The gelation bath is a blend of 70 to 60% acetone in water.J). polyvinylidene difluoride. They include the use of additives in the casting dope or in the precipitant (water) and temperature adjustments to both.9’ The maximum water content. For example. polycarbonate. Mg(ClO& on the porosity of CA membranes. Since very little acetone has diffused out of the casting solution. polymethyl methacrylate. This is because additives in the casting dope would be expected to reduce the chemical potential of other species in the casting dope. Reducing the chemical potential of water in the dope will increase the chemical potential driving force for water entry. polymer begins to precipitate around a large volume of solvent resulting in a very porous membrane. The result is an MF membrane with relatively large pores (over 0. Eirich investigated cellulose acetate. polyacrylonitrile. Formamide is a standard additive for CA-acetone casting solutions. Various methods have been used to adjust the rate of water entry and solvent removal. and temperature. This produces tighter membranes with lower porosity.146
Handbook of Industrial Membrane Technology
bounds. At this level. On the other hand. the nature of the solvent can have a marked influence on the transport properties of the resulting membrane. it is well known that an increase in formamide concentration will increase membrane permeability and decrease salt rejection. polyvinyl chloride. if the additive in the water bath is the same solvent used in the casting solution. Again. decreasing the rate of water entry into the cast film because of the decreased chemical potential of water. One may also reduce the chemical potential driving force for water entry by reducing the temperature of the water bath. Figure 3.6) in the formation of RO membranes. reducing the chemical potential of solvent in the casting dope will decrease the chemical potential driving force for solvent removal. If the solids are increased by direct addition to the casting solution or by an evaporation step (moving from point A to A’ on Figure 3. polystyrene.8 The casting solution is 16 to 19% PVDF in acetone at a temperature of 5O’C. this is readily explained by the ternary phase diagram shown in Figure 3. is only slightly above 15%. Likewise.1 /. the highest porosity membranes were made from casting solutions utilizing solvents with high solubility parameters. any additive in the casting dope would be expected to increase the rate of water entry and decrease the rate of solvent removal-resulting in more open membranes. If the acetone concentration in the water bath is reduced below 59%. only a small amount of water need diffuse in before gelation occurs. a skin begins to form and an asymmetric UF membrane is produced. the rate of solvent removal from the cast film is also slowed along with the rate of water entry.7). Additives in the water bath can have the opposite effect. The result can be a reasonably homogeneous pore structure with fairly large pores. From a thermodynamic standpoint.

5 were dimethylformamide casting solutions which gave membranes of lower porosity than the correlation would predict. the trend was the same-the water content (porosity) increased with higher solubility parameters of the solvent.
Since water (the precipitant) has a solubility parameter (23. Ethyl formote Propylene oxide
1 10 Solvent solubility
1 11 porometer
I 12
Figure 3. Likewise.
.4) higher than any of the solvents investigated. since a polymer’s solubility parameter cannot be calculated because of its nonvolatility. The data for cellulose acetate are plotted in Figure 3. only 7 cases were out of line. Of the 7 exceptions. two solvents with similar solubility parameters are more compatible and have a greater affinity for each other.109 Though some of the polymers investigated were not soluble in all of the solvents shown in Figure 3. The solubility parameter for solvents is calculated as the square root of the cohesive energy density (the energy of vaporization divided by the molar volume). In general. hydrogen bonding or dipolar interactions. Of 41 polymer-solvent systems investigated.
n-methyl
Dimethyl ocetomide pyrrolidone
Dimethy
sulfoxidf
230I”
’
‘“I 10 0 0%
I
.148
Handbook of Industrial Membrane Technology
chloride copolymer with a whole host of solvents. whether they be dispersion (London) forces.10. Tetrahydrofuran always resulted in very dense membranes while dimethyl sulfoxide yielded very porous membranes. it would appear experimentally that the disparity between the solubility parameters of the solvent and the precipitant is a key factor in understanding these results.10: CA membrane porosity (water content) as a function of the solvent used in the casting dope. It is a measure of all cohesive forces tending to hold the molecules of the solvent together.

commercially available fibers cannot withstand internal pressures up to 400 psi and above.Ultrafiltration
149
it is normally determined experimentally by arranging a list of solvents for the polymer in order of increasing solubility parameter.
Figure 3. low porosity membrane. smaller differences in the solubility parameter would tend to favor inclusion of water resulting in a membrane with higher porosity.
. This is because of the high pressure required in RO. Procedure for Spinning Hollow Fibers
Figure 3. On the other hand. The result is a dense. Thus. The simplest explanation for the data of Eirich (though not suggested by Eirich or his co-workers) seems to be related to the greater compatibility of water with solvents having higher solubility parameters. the solubility parameter for the polymer is taken to be that at the midpoint of the soluble range. as we shall see. and the permeate flows from the lumen of the fiber (see Figure 3.11 is a photomicrograph of a typical asymmetric hollow fiber with the skin on the inside wall.12) . When the polymer begins to precipitate. Hollow fibers for UF have the skin on the inside whereas hollow fibers for RO are smaller and have the skin on the outside. for RO.Actually.11:
Photomicrograph
of asymmetric
U F hollow fiber . the feed stream is pressurized on the outside of the fiber. flowing the feed stream down the lumen of the fiber greatly reduces concentration polarization effects. the residual solvent may tend to exclude water if there is too great a disparity between the solubility parameter. considerable research effort has been spent in developing composite hollow fibers for RO which can withstand internal pressures up to 900 psi because.

14 shows UF fibers made by Romicon with inside fiber diameters up to 1 mm (1.
The lower pressures required in UF make possible better fluid management with pressurized feed on the inside of the fiber (see Figure 3. with feed in the bore of the fiber. Naturally.
Figure 3. the inside diameter must be larger to prevent plugging. Since most UF hollow fibers have burst pressures ranging between 60 and 100 psi. the trend is to go with even larger diameters (1.13).5 and 2.
. inlet operating pressures are limited to an inlet pressure of 30 psi.13:
Feed stream on inside of UF hollow fiber. Figure 3.000 ~1.150
Handbook of Industrial Membrane Technology Hollow fiber principle
Figure 3.D. Actually.0 mm I.D.12:
Feed stream on outside of RO hollow fiber. This is 25 times larger than the inside diameter of a DuPont RO hollow fiber (42 JL I.).) to improve fouling resistance.

a secondary skin will form on the outside of the fiber (see Figure 3. The degassed and filtered polymer solution is forced under pressure into a coaxial tube spinneret. so that the outside of the emerging fiber is in contact with air before being collected in the water bath.15:
Schematic of tube-in-orifice spinneret. The liquid is extruded through an annular orifice and the hollow fiber (still liquid) is stabilized and precipitated by an internal coagulating fluid (usually water) which flows out the center tube.14: Dimensions of various hollow fibers supplied by Romicon (UF) and by DuPont (RO).
.
The spinneret is usually positioned some distance above a water bath.
The spinning of asymmetric hollow fibers with the skin on the inside closely resembles the procedure used in casting flat-sheet membranes.76).Ultrafiltration
151
Figure 3.
Figure 3.151° is a schematic diagram of a spinneret used to spin these fibers. If the fiber hits the water bath before being fully gelled. Figure 3.

Spinning Variables for Hollow Fibers
All of the variables discussed in conjunction with casting sheet and tubes apply to hollow fibers as well.16: fiber. etc. and reduced temperature of the coagulant will result in tighter. Higher polymer content in the casting dope. this water bath is continuously replenished with fresh water. additives in the casting dope and solvents with high solubility parameters favor more open fibers with a higher permeability.
Photomicrograph
showing secondary skin on outside of UF hollow
This is disadvantageous when potting the hollow fibers in a header.
. The spinneret design (the diameter of the center tube and the width of the annular space between the inner tube and the extrusion orifice) will determine the dimensions of the final fiber. to prepare for drying before potting. there are other variables unique to the spinning process which offer additional handles for controlling fiber permeability and retentivity. The spinning procedure accomplishes each of the six steps used in casting flat sheets or tubes.152
Handbook of Industrial
Membrane Technology
Figure 3. In addition to the above. additives in the coagulating fluid. Collection of the gelled fiber in a water bath permits leaching of remaining solvent from the fiber. The potting material (usually epoxy) will not penetrate into the wall of the fiber resulting in a poor bond and a possible leak path for the feed stream through the cross section of the fiber wall and out into the permeate. During long runs. the fibers may b~ collected and immersed in other baths to add plasticizers. by-passing the internal skin. On the other hand. If need be. more retentive hollow fibers.

rectangular pores of the same area as circular pores will show retention for smaller molecules and a lower permeability. the permeability through the rectangular pores will be only 2/3 of that through the circular pores. As a consequence. an increase in the velocity of the coagulating fluid through the inner tube tends to counteract the elongation of pores in the skin. very small “air-gaps” can result in the formation of a secondary skin on the outer surface of the fiber which may tend to decrease the hydraulic permeability as well-though the retentivity of the fiber (determined by the internal skin) would be more open. Likewise. However.
153
In general. An increase in the coagulating fluid velocity will have the opposite effect. this alignment of polymer chains will be frozen in place in the solid skin. leading to more open fibers of higher permeability and reduced retentivity. If the kinetics of gelation remain unchanged. It is known that polymer molecules under shear flow tend to align themselves in the direction of flow. if the width of the rectangular pore is equal to the radius of the circular pore of the same area. Since the internal wall of the fiber precipitates immediately upon contact with the coagulant after emerging from the spinneret. the increased pressure in the fiber lumen tendstoexpand itsdiameter and thereby enlarge the width of the elongated pores. This inorganic dynamic membrane could have flux values as high as 1. While numerous colloidal materials and organic polyelectrolytes were tested. will further extend the polymer chains as they passthrough the spinneret because of the added weight of the fiber below. an increase in the “air-gap” (distance between the spinneret and the collecting bath). In the late sixties.000 GSFD (gallons per square foot per day) at 950 psi with a NaCl rejec-
. (3) The distance between the spinneret and the collecting bath (the “air-gap”). Preparation of inorganic Membranes Dynamic Membranes. (2) The velocity of the coagulating fluid through the inner tube. Indeed. Thus.Ultrafiltration (1) The velocity of the spinning solution through the annular orifice of the spinneret. the fiber. The effect of increased spinning solution velocity in producing tighter fibers is explained by the shear forces in the annular orifice. On the other hand. Any tendency towards relaxation of the polymer chains after emerging from the spinneret and before gelation will be alleviated by this weight tending to stretch. the most successful system was based on hydrous zirconium oxide. The alignment of polymer chains in shear flow through the spinneret may explain why hollow fibers generally have lower permeabilities than flat-sheet membranes with the same retention rating. The result is a more open fiber with increased permeability and decreased solute rejection. workers” at the Oak Ridge National Laboratory discovered that cross-flow filtration of dilute colloids through a microporous tube made the support tube retentive for macromolecules and even salts (in some cases). an increase in the spinning solution velocity and the “air-gap” will result in tighter fibers with lower permeability and increased retentivity. the pores in the skin (which control retention and permeability) will be elongated. an increase in the “air-gap” will only increase the degree of orientation resulting in tighter fibers.

This resulted in a high-temperature (300°F) inorganic membrane which is resistant to a pH of I-14. it is stripped off with acid and a new membrane applied using cross-flow of the Zr( IV) oxide slurry feed. MA) is reportedly about to introduce an alumina UF membrane in addition to its current MF membrane.13 They use a sol/gel technique with Boehmite (y-AIO-OH) as the precursor because it can be easily dispersed with acids. This technology was subsequently licensed to SFEC (Soci&i de Fabrication d’El&ments Catalytiques) in Boll&e. the membrane may be cleaned with organicsolvents that would dissolve conventional polymer membranes. NC. France and to Gaston County Filtration Systems in Stanley. Further. one method for preparing gamma-alumina membrane films on porous supports (which can be made from alpha-alumina) has been reported by Dutch researchers. Rejections of 90% for 0. The difficulty with existing ceramic UF membranes is resistance to high pH. Alumina membranes for UF have been introduced by Ceraver (Tarbes. SC) has exploited this technology by selling plants utilizing porous stainless steel tubes and the know-how to apply the hydrous Zr(lV) oxide dynamic membrane. The alpha form of A120J is used to make MF membranes (0. The Norton Co.05 M). Seneca. Ceramic Membranes. Union Carbide went one step further-sintering the hydrous Zr(lV) oxide dynamic membrane in place on a porous carbon tube at elevated temperatures.05 M NaCl were obtained. (PAA) at high pH on top of the Zr(lV) oxide dynamic membrane. France) (membrane division recently purchased by Alcoa).12 One company (Carre Inc. >8O=‘C [ C&CH
tCH$O$
Al
+
Ws&r
y-AIO-OH
Boehmi te (preclpl tate)
ACM
) 90-c
ealciMtlal FiJ pore size Y A& film Thermal Treatment t (Temp/tlme) t-10 mlcma W
Y A&
+
) 4OO’C
y -AIO-OH
(sol)
. the existing technology uses the gamma form of A1203 which will not withstand high pH.154
Handbook of Industrial Membrane Technology
tion of 50% at dilute salt concentrations (0.2 /. (Worcester..r pore size and above) and is resistant to a pH of I-14. When the membrane becomes fouled. However. Further. the rejection could be improved by applying a thin layer of polyacrylic acid. The methods used to make these membranes are diverse and highly proprietary. To make the finer pores (down to 40 A) in a UF membrane.

Unfortunately. The convention established by AMICON and adopted by most UF membrane manufacturers is based on the retention of globular proteins (spherical macromolecules). It is difficult to measure the pores directly by any of the techniques used for MF membranes.
1. The retention R (in percent) may be defined as follows:
(2)
where
= =
R=lOO CR lC”f CR
( )
cuf
concentration of the solute in the ultrafiltrate concentration of the solute in the retentate
. The porous support is dipped into the sol and capillary action pulls tlie sol into the pores increasing the concentration of Boehmite at the entrance of the pores to form a gel. UF membrane pore sizes range from 10 to 1000 A (0. Most UF membranes have “diffuse cut-off” characteristics. we are accustomed to think of the molecular weight of the macromolecules which are retained by or pass through the membrane.03 to 0.17:
Sharp vs diffuse cut-off membranes.17). The calcination of this gel film (above 4OO’C) yields the final film of gamma-alumina. Instead.2 (in the Preface).000 wt.Ultrafiltration
155
The pore size of the membrane is determined by the type of acid and its concentration during dispersion (typically 0.000 100.1 II).001 to 0. The anisotropic structure and the wide distribution of pore sizes make this almost impossible.15 mol acid per mol of Al&OH) along with the final thermal treatment. Log molecular Figure 3.
Molecular Weight Cut-Off and Pore Size What we mean by the molecular weight cut-off of a “diffuse cut-off” membrane must be clearly defined. both of which affect the crystallite size of the gamma-alumina. existing UF membranes do not have a “sharp cut-off” because of the wide distribution of pores in the skin of the membrane (see Figure 3.
PORE SIZE DETERMINATION As shown in Figure P.0 -
Retention
0’ 1000 10.

Figure 3.16 shows how this is determined.
. The molecular weight at which the retention curve crosses a retentivity of 90% is the “molecular weight cut-off” of the membrane. For example.19:
Pore size variation with MWCO. The retention values of a series of globular proteins (spherical molecules) are measured on the same membrane. This may be used to estimate the effective pore size (see Figure 3.18: branes.
IO IS APPROXIMATE 20 MOLECULAR 30 40 DIAMETER so 60 ifi1 SO 100
Isa
X.000 molecular weight cut-off (MWCOI.
Retentivity
of a series of globular proteins on various UF mem-
Id
MOLLCLMAR WEIOHI
Id
CUT-OFF (MWCO)
Figure 3. This means that larger molecules are said to be retained by the membrane and smaller molecules are said to pass.
5
IO’
I
2
.18. The approximate molecular diameter @I at the top of Figure 3.156
Handbook of Industrial Membrane Technology
The convention states that the molecular weight cut-off of the membrane is equal to the molecular weight of globular proteins which are 90% retained by the membrane. UM. PM-10 indicates the membrane has a 10. in Figure 3.
5
MOLECULAR
WEIGHT
Figure 3.18 is taken from radius of gyration data reported in the literature.19).NOMINAL
MOLECULAR 5 IO3
WEIGHT 2
EXCLUSION CL I04
OF MEMBRANE 2
01
I02
2
*
I
*
I
. Normally UF membranes are designated by prefix letters which refer to the membrane type followed by one or more digits which refer to the molecular weight cut-off in thousands of daltons. PM and XM refer to three different polymers.

Flexible Polymers
on UF Retentivity
Membrane* Diaflo XMSO Diiflo PM30
Dextran 250 (236.s Globular Proteins T-Globulin (160. Each membrane retains those molecules positioned above the line but passes those below the line.000) Cvtochrorne C (13.000) Pepsin (3S.
Figure 3. Apparently.000) Polyacrylic acid (pH 10.000) Insulin (5.21 which shows the difference in the retentivity of a 1.114 shows the effect of size and shape of the molecule on its retentivity by various MWCO membranes. the retention of branched polysaccharides and linear flexible proteins is lower than might be expected. the fluid shear in the vicinity of the pores is high enough to uncoil free draining chains. However. Polyethylene glyc01(20.20:
Retention
of spherical
and linear molecules. but of greater diameter. Dextran 250 with a molecular weight of 236.000) Branched Polysaccharides Linear.000) Diaflo PM10 Bacitracin (1.400) Dextran 40 (40.20). a linear polymer may snake through the pores while a spherical polymer of the same molecular weight. the retention of these molecules is as expected.000)
. ** Number in parentheses denote molecular weights. Since characterization of a membrane’s MWCO is based on globular proteins.000) Polyacrylic acid (pH 7.Ultrafiltration Retention of Spherical and Linear Molecules
157
Two different molecules of the same molecular weight can have different configurations such that their molecular diameter is different. is retained (see Figure 3.000)” Albumin (69.000 passes through a 50. For example.000) Diaflo UMIO * Molecules above J horizontal membmne line are completely retained by the membrane. For example. Table 3.000) Dextran 10 (10. 50. 50.700) Dextran 1 IO (100. below the line partial retention or complete clearance is observed.000 MWCO membrane. This is illustrated in Figure 3.000 MWCO membrane for a series of globular proteins and a comparable series of polysaccharides.1: Effect of Size and Shape of Molecules
Solute C1x.000.
Table 3.

000 MWCO membrane can have a (I.
.21: membrane.OOO.oac. small laboratory discs of these membranes can be subjected to high challenge levels of bacteria with absolute retention (zero passage). Figure 3. it is difficult to manufacture a pinhole-free module with this much area.P.) bubble point of 100 psi.OOO MWCO
It is also known that the pH and ionic strength of solutions of polyelectrolytes can have a marked influence on their retentivities.158
Handbook of Industrial Membrane Technology
looo MEMBRANE
0. Equation 3 of Chapter 2 may be used to calculate a maximum pore diameter of 0. the larger the effective size of the polyelectrolyte for a given molecular weight. Broken fibers.OOO.OOO.yet the bubble point indicates a maximum pore size in the skin over 0. it is obvious that larger pores do exist. and the lower the ionic strength of the medium.22 also indicates that a 10. However. Maximum Pore Size and Effective Pore Size Since the effective pore size is estimated from the molecular diameter of globular proteins which are retained 90% by the membrane. and other defects provide leak paths for bacteria.000 MWCO membrane (F300) should have an estimated effective pore size of 0. 1.r which should be retentive for all bacteria.02 /_I.10 4 5
MOLECULAS WElbW 6789. The more a polyelectrolyte is charged in solution. 2
Figure 3.000 and 1 . However.
Retentivity of proteins and polysaccharides on a 1.4 /.06o. bubbles in glue-line seals.A 300.12 /. The measurement of a membrane’s bubble point (see the section on the bubble point test in Chapter 2) permits calculation of the maximum pore size in the skin of the membrane. Indeed. industrial scale UF modules often employ 10 to 100 square feet of membrane area.L This is one reason why UF membranes can be less retentive for bacteria than MF membranes.A. Figure 3.22 shows the bubble point measured with isopropanol (IPA) on polyvinylidene difluoride UF membranes with MWCO’s between 10.

22:
IPA bubble point of a series of UF membranes.0s
C
Figure 3.) (3)
R = loo
ldvo/vr 1
Equation 3 is used to calculate retention solely on the basis of changes in the retentate volume and concentration. Adsorption Losses The retention of membranes are often measured in stirred cells.0.Wi
6.
RETENTION
CHARACTERISTICS
The polymer from which a UF membrane is made does not generally affect the retention characteristics of the membrane. when processing small volumes of dilute solutions. and this can affect the apparent retention calculated from Equation 2. the retention calculated from Equation 3 will be too low since Cf will be lower than it would be without adsorption. However. For this reason.h 4 tz
\
p&T
f36
i
_ T\ FYY) t_9++ F= F5ao
(RtlCROh.
%RE
s\%bE
I
I
lllll
I
I
I
IllIl_
lMo2
O. if adsorption losses are ap preciable.62
403
0. For example.+-)
/
EST1 M n= I $0. integrated over time from an initial retentate volume (V. A mass balance on the cell. strong adsorption of the solute can diminish its concentration in both the retentate and in the filtrate. A simple cumulative mass balance on the cell leads to another expression for
. the nature of the polymer can affect adsorption of species onto the membrane surface. Obviously.) and initial concentration (Co) to a final retentate volume (Vf) and final retentate concentration (cf) yields the following expression for the retention (R): ln ( cr / C.01
0. the retention should also be calculated with reference to the solute concentration in the ultrafiltrate (permeate)..Ultrafiltration
159
100 -f
!302#
Q-5 8
‘El
I
FIO 0
BUBSE SERIES
POINT
OF R
_ 1 _
C4= ULTRRFILTRRTIM
hEH6RRNR
(ISOPROPYLALC4HOL)
I
Fzo
‘lo-1 -3 2.

This technique may be used to measure the relative adsorption of various membranes as in Table 3. adsorption losses can be significant.3 4. If a small volume of a dilute solution is placed in contact with a large membrane area. the initial concentration of the retentate (C. the retentivities calculated from Equations 3 and 4 may be compared.3 11.IO
k4MW8t9d
c8lIuless AceWe
&hJloSe Acetate
0.160
Handbook of Industrial Membrane Technology
retention (R) based on the final (average) ultra filtrate concentration (Cut). If there is an appreciable difference.6
97 97 26 a99 85 >99 53
Amimn YM-10 Amimn UN.). Fifty ml of a dilute solution (150 micrograms/ml) of cytochrome “C” were concentrated to 10 ml in a 50 ml stirred cell (membrane area of 13. If the initial and final retentate concentration (C.2. Table 3.5 24. hydrophilic polymers tend to have a lower adsorption than the hydrophobic polymers.) and charges in the retentate volume from V. to Vf:
R=
Again.2. the retention calculated from Equation 4 will be too high since C.
.4 22. if there are significant adsorption losses. The magnitude of the loss may be calculated from the amount of solute in the retentate (CfVf) and in the filtrate K&D/o-Vfl 1 compared with that charged to the cell (CoV. and Cf) are measured along with the final ultrafiltrate concentration (Cur).2: Adsorption Lossesof Cytochrome C on Various UF Membranes Membrane Types Adsorption Retention for CytochromeC(W1
loss(%)
Nudepore C.f will be lower than it would be without adsorption.10 Millipore PTOC I 04) ( Nuclepore IO ANuclepore IO FAmlmnPM-IO
@JfW8td
Polyelactrolyte Complex Polysulfons Polyamide Polyvinylkbnedifluerids Polyarylsulfens
In general..3 12. the adsorption from solution can be appreciable.8 2.4 cm’) to compare the various membranes listed in Table 3.

If “pore-flow” were in effect. leakage of salt through large pores would also be expected to increase with pressure. the increase in salt rejection with increasing pressure suggests a “solution-diffusion” model as opposed to a “pore-flow” model for transport of solvent and solute across the membrane. In the case of RO (Figure 3. Highly hydrated ions are rejected better than poorly hydrated ions. divalent and trivalent ions are rejected better than monovalent ions. It is known that although the water flux through RO membranes increases with pressure. Effect of Pressure Figure 3. psi Figure 3. Charged UF membranes reject low concentrations of salts primarily by the Donnan exclusion mechanism.1 M NaCl by a cellulose acetate RO membrane.‘St16. the Millipore PSAL membranes contain fixed sulfonic acid groups which have a negative charge capacity of 800 millimols/liter.
.
FIGURES
IN BRACKETS
INDICATE
ORDER OF RUNS
I I I 1 40 20 60 80 TRANSMEMBRANE PRESSURE DROP.2317 presents rejection data for three different dextrans on the same UF membrane as a function of pressure.24ls shows the effect of pressure on the rejection of 0. It is obvious that the mechanism for solute transport through the membrane is different for UF and RO.24). Because the fixed charged groups on the membrane skin reject ionic solutes via repulsion of coions.For example. the salt flux is almost invariant with pressure. This suggests that the RO membrane acts as a nonporous diffusion barrier.23: Effect of pressure on UF rejection of dextrans. the rejection would be expected to depend on solute type and coion charge.Ultrafiltration “Charged” Membranes
161
Polyelectrolyte complex UF membranes can be made with a net charge to obtain moderate rejection of salts or amino acids (75 to 90%) at relatively high flux and low pressure. Figure 3. The net result is that a higher water flux dilutes the salt concentration in the filtrate resulting in a higher calculated rejection by Equation 2. Obviously.

For water. The water flux is higher than the salt flux because the soiubiiity of water in the membrane and its diffusivity through the membrane is higher than that of salt. there is little difference in concentration across the membrane. For salt.24: Effect of pressure on RO rejection of 0.
.1 M NaCi. This means the dichiorophenoi passesthrough the membrane more readily than water. This is hard to explain with a “pore-flow” sieving mechanism.) should be proportional to the water flux (J. the rate of transport of water and salt is proportional to the chemical potential gradient of each species across the membrane. For example.) which carries solute to the membrane and through the larger pores. the dominant effect is the difference in pressure. cellulose acetate membranes which show a 95% rejection for NaCi (MW 58) and a 99% rejection for dextrose (MW 180). the solute flux (J. The “solution-diffusion” model explains why molecules larger than salt sometimes pass through an RO membrane more readily. show a negative -34% rejection for 2. in a “pore-flow process”.
in a “solution-diffusion” process. the ratio of upstream to downstream salt concentration is very large and this dominates over the effect of the pressure difference. The chemical potential of each is affected by pressure and concentration.4-dichlorophenol (MW 163).162
Handbook of industrial Membrane Technology 100
90 -
l
0
60 l
increasing pressure Decreasing pressure
I I I
50
I
I
I
I
I
0
10
20
30
40
50
60
70
80
90
Applied pressure. atm Figure 3.

both fluxes increase.(dc/dx).) has increased over that in the bulk stream (Cc).
Thus. there must be a concentration gradient (dc/dx) to remove solute from the membrane.23. the phenomenon of “concentration polarization” may cause the rejection to decrease with increasing pressure as in Figure 3. carrying a higher concentration of solute (Cs).Ultrafiltration
163
(5) where Cs c = = concentration membrane of
Js = u J&s
solute at the upstream surface of the
the fraction of the total liquid enough to pass solute molecules for solute may be expressed
JS
flowing
through
pores large
The retention
as:
(6)
or (7)
R=
I-
J
W
R= 1 UC. As the pressure is increased. The reason is that the concentration of solute at the surface of the membrane (C. the solvent flux is increased and the convective transport of solute (J. pass more solvent.23. it accumulates at the surface of the membrane until the back diffusive mass transport. at higher pressures.25:
Concentration
polarization
in UF. Thus. (=J. R would not be expected to increase with pressure since the solute and solvent flux are “coupled”. D.C) to the membrane is increased (see Figure 3. However. Even if a steady state is maintained.
. in a “pore-flow” model. The pores large enough to pass solute. is equal to the forward convective transport.25). Cs will increase with pressure resulting in a decrease in solute retention by Equation 7 and as shown in Figure 3.
‘b
J. As the pressure is increased.C)
Figure 3. If the solute is retained by the membrane.

26).) between 20 and 50 wt %. The retention of individual solutes from a mixture is more complicated. Fractionation may or may not be possible. Indeed. Because the gamma globulin is retained by the membrane. and adsorption characteristics) and by the operating conditions (e. pressure. the membrane is said to be “gel-polarized“. Under these conditions (see Figure 3.164
Handbook of Industrial Membrane Technology
Further increases in pressure will increase the solute concentration at the surface of the membrane (C. the rejection of Dextran-80 is beginning to level out at the higher pressures as the membrane becomes gel-polarized. charge.000 daltons) of 95% and a retention for albumin (67. the retention should be independent of pressure (see Equation 7).
Fractionation of Solutes The retention of a single solute is determined solely by the properties of the membrane (pore size distribution.) to a limiting concentration. the rejection of albumin approaches the single solute rejection by the membrane. Colloidal suspensions form a densely packed layer of closepacked spheres usually between 70 and 80 volume %. The retention for albumin in the presence of gamma globulin is shown in Figure 3.23. the albumin would pass the primary membrane. at low concentrations of gamma globulin. pH and ionic strength of the solution). In Figure 3.
. once the membrane is “gelpolarized”. Consider a 100.000 MWCO membrane which has a retention for gamma globulin (160. Therefore.. Without this “dynamic secondary membrane”.
MEMBRANE
LAMINAR \ “CAKE” SUBLAYER (NON NEih’TONIAN)
Figure 3.g.000 daltons) of less than 15%..27. it forms a gel-layer which is retentive for albumin.26:
(SLIME)
Gel formation due to concentration polarization. Proteins begin to form a semisolid gel on the membrane surface with a gel concentration (C. and further increases in pressure will not increase C.

2% CYTOCHROME C PM-30 MEMBRANE STIRRED CELL (5-55 psi) IO
I 1.000
8 I g K 0 LL o z F:
MW) (26.000 MW)
I
MW)
50-
3 30t.1 0. ez. they would have to dilute the mixture way down to accomplish the separation. 200
0
W
A OVALEUMIN (45.27:
Retention
of albumin
in the presence of gamma
globulin.2% solution of cytochrome C (12.0 CONCENTRATION OF MACROSOLUT: of cytochrome C in the presence of larger species. Figure 3.000 0 CHYMOTRYPSINOGEN
0 100 ALBUMIN (68. lwz $ ti IO ’ 0.
The pharmaceutical industry would prefer to make the gamma globulin/albumin separation with a UF membrane rather than Cohn fractionation (sequential precipitation with ethanol).
Figure 3.Ultrafiltration
165
Figure 3.000 MWCO membrane influence the retention of a 0.400 daltons) which by itself passes through the membrane with 0% retention.28 shows how three different solutes which are retained by a 30.28:
. But to do so. Separations involving globular proteins appear to be the most difficult.3 Retention
SYSTEM:
l/2X UF 0. The processing of the large diluted volumes followed by concentration of the two fractions would make the membrane process more cumbersome and expensive than Cohn fractionation.

thereby reducing the flux.. Without the development of an anisotropic UF membrane. polymers like hydroxyethyl starch.29)” can be fractionated more readily. to that in the bulk process stream. Consequently. it will be seen that concentration polarization can severely limit the flux.
MEMBRANE
FLUX WITH CONCENTRATION
POLARIZATION
The effect of “concentration polarization” on retention was discussed in the previous section. The control of polarization by proper fluid management techniques is essential to the economic feasibility of the process.25 can severely limit the flux. Cb) is higher than that experienced with lower permeability RO membranes. C. the salts retained have a significant osmotic pressure (a) and the effective pressure gradient is reduced by the osmotic pressure difference across the membrane (An).20
WEIGHT DISTRIBUTION OF DIAFILTERED POLYVINYLPYRROLIDONE DIAFLO XM-48 MEMBRANE
I
I
I
I
I
RElENTAlE
mi
2ooJJool
I
Figure 3. In this section. MOLECULAR
0. However. polyethylene glycol.29:
Fractionation
of PVP. In the case of RO. The thin skin minimizes the resistance to flow. UF would not be a commercial process today. it is argued that the high solute
.20-22 Even though the macromolecules and colloidal suspensions retained by UF membranes are quite large and have negligible osmotic pressures.166
Handbook
of Industrial
Membrane
Technology
On the other hand. This accumulation of solute at the membrane interface (see Figure 3. the polarization modulus (defined as the ratio of the solute concentration at the membrane surface. and polyvinyl pyrrolidone (see Figure 3. and the asymmetry of the pores virtually eliminates internal pore fouling. Some researchers have used this “osmotic pressure model” in an attempt to explain the effect of concentration polarization on UF membrane flux. the hydraulic permeability of these membranes also increases the convective ransport of solutes to the membrane surface.

This “secondary membrane” can offer the major resistance to flow. If the system is not “gel-polarized”. One of the curious aspects of data like those in Figure 3. Effect of Pressure. by the very nature of the process. Gel-Polarization If the transmembrane pressure drop (Ap) and the solute concentration in the bulk process stream (Cb) are high enough.31. as in Figure 3. the concentration at the membrane surface (C. the flux in stagnant “dead-ended” systems is often so small as to be virtually nonexistent unless the bulk stream concentration is extremely low.) increases resulting in an increase in the concentration driven back diffusive transport away from the membrane. when flux is plotted versus pressure. Thus. Figure 3. and pressure.26). the gel concentration at the membrane surface is fixed (C. Cs will increase until the backdiffusive transport of solute just equals the forward convective transport. can be calculated from the experimental flux. increased pressures will not help since the gel layer only grows thicker to offer more resistance to the increased driving force.Ultrafiltration
167
concentrations at the membrane surface can result in an osmotic pressure difference across the membrane which should be taken into account.). At present.25). The increase results in a higher rate of convective transport of solute to surface of the membrane. is a constant which can be calculated from the pure water flux. In a stagnant “dead-ended” system.) as a function of stirrer speed (in a stirred cell). the following treatment will neglect the osmotic pressure. the flux does increase initially. the flux often becomes independent of pressure in the steady state.30 is that flux does not increase monotonically with pressure.30 is a plot of the cake resistance (R. Therefore. the gel layer will grow in thickness until the pressure activated convective transport of solute with solvent towards the membrane surface just equals the concentration gradient activated diffusive transport away from the surface. Lower protein concentrations in the bulk (Cb) also increase the concentration gradient.
(8)
where
Jw=_
‘1
Jw = water flux (volume/time/membrane AP = transmembrane pressure drop
area)
Rm = hydraulic resistance of the membrane Rc = hydraulic resistance of the deposited cake Since R. Furthermore. In fact.
. R.) can rise to the point of incipient gel precipitation forming a dynamic “secondary membrane” on top of the primary structure (see Figure 3. the boundary layer thickness decreases (see Figure 3. protein concentration. thereby increasing theconcentration gradient (dc/dx) for removal of the cake. the solute concentration at the surface (C. the “gel-polarization model” appears to do a better job predicting the UF flux for a wider range of process streams than the “osmotic-pressure model”. Indeed. When the pressure is increased. As the stirrer speed increases.

exp(-a = diffusivity at infinite dilution = constant
Cl
If one integrates Equation 9 taking into account Equation 12. Further increases in pressure will again temporarily increase the convective transport (J. in the “gel-polarized regime”. It is recognized that the diffusivity (D) is really a function of the concentration profile near the membrane surface. However. and solute will accumulate on the membrane. the convective transport to the membrane must equal the back-diffusive transport away from the membrane.) to the membrane surface. since the surface concentration is at a maximum. The net result is that the flux will decrease to its original value in the steady state. The gel layer will thicken or compact just compensating for the increased driving force (API by an equal increase in the resistance of the cake (R. In the steady state. flux is independent of pressure and is solely determined by the back-diffusive transport. the final equation is considerably more cumbersome than Equation 10 without much gain
. at the membrane surface. = C. the back diffusive transport will be fixed (assuming no changes in the fluid dynamics in the boundary layer). the boundary conditions are fixed: C C = Cb at large distances away from the membrane = C.).). Therefore.
Integrating Equation 9 assuming that the diffusivity (D) is constant
(IO)
J
where K is the mass transfer coefficient: (11) KD ?7
and 6 is the boundary layer thickness. (II
D = D.Ultrafiltration
169
Eventually the concentration at the membrane surface will be high enough for a gel to form (C. dc JC= where J C D
X
-Ds
= solvent flux through the membrane = concentration of membrane retained solutes or colloids = solute diffusivity
=
distance from the membrane surface
In the gel-polarized regime. If one accepts an exponential mode of the diffusivity:
(12)
where D.

at the same concentration.
0. higher stirrer speeds result in a higher asymptotic flux. because the boundary layer thickness (6) has been reducedincreasing the mass transport coefficient (K) (see Equations 10 and 11).8. Once the gel-layer is formed. I 5 8 I 5
IO. Equation 10 applies to all forms of fluid management-cross-flow as well as stirred cells. Lower solute concentrations (Cb) have higherthreshold pressures (Pt). yet in the presence of protein (retained by both).7. the concentration gradient (C. it is often the limiting resistance to flow. Figure 3.).5 *
. It will be noted that the higher hydraulic permeability of PM 30 membrane results in a much lower threshold pressure (7 psi)..32 shows two membranes with widely different membrane resistances (I$.
Of course.32:
Limiting resistance of gel layer
membrane resistance).0
lRANSM&BRANE
PRESSURE
(psi) (VS
Figure 3. cross-flow techniques are preferred-where the feed stream flows tangential to the membrane surface (see
.31 shows that below some “threshold pressure” (Pr).170
Handbook of Industrial Membrane Technology
in accuracy. This also means that the threshold pressure (Pt) will be higher for higher stirrer speeds.31 is higher for lower protein concentrations. Likewise.I 1. The higher removal rate of solute from the membrane requires higher pressures and flux to carry enough solute to the membrane to form a gel. The pure water flux differs by a factor of 3.030
I . This also means that the asymptotic flux in the gel-polarized region of Figure 3. Low solute concentrations favor a high back-diffusive transport of solute. For large scale systems. and a much higher flux is required to transport enough solute to the membrane (JC) to begin to form a gel. Figure 3. Equation 10 shows that the flux is independent of pressure in the gel polarized regime..75. for pressures over the threshold pressure of 20 psi. flux still increases with pressure.73
A
.23 Suffice it to say that the calculation of flux from Equation 10 will overestimate the flux slightly because the diffusivity in the boundary layer is lower than that at infinite dilution (De). the water flux differs by a factor of only 1 .-Cb) driving back-diffusive transport is higher.

Kin (C.) . a restrictor is placed on the exit retentate stream to keep the exit pressure high so as to maximize flux throughout the channel length. the equivalent of a gel layer should be a layer of close packed spheres having a packing density between 60 and 75%. Figure 3.33). Some data in the literature show flux vs.34 are entirely reasonable.Ultrafiltration
171
Figure 3.34 is equal to the mass transfer coefficient which is a strong function of the stirrer speed or the tangential velocity across the membrane.j
The form of Equation 13 has been demonstrated for a large number of macromolecular solutions and colloidal suspensions. which is proportional to the tangential velocity across the membrane). = 28%). the intercept with the abscissa should occur when Cb = C.) For laminar flow. There is a pressure drop down the channel or tube which means that the first part of the channel may be gel-polarized. Likewise. All lines converge at zero flux where the concentration equals the gel concentration (C. the viscosity increases with concentration requiring constant adjustment of the pump. Many protein solutions reach their solubility limit between 25 and 35%. The negative slope of each of the straight lines in Figure 3. Often the reason for the curvature is the inconstancy of the recirculation rate. the entrance region of the channel may not be gel-polarized either.33: Cross-flow UF system. If Equation 10 applies. for colloidal suspensions. as the
. it should be possible to plot the solvent flux versus the logarithm of the concentration (Cb) and get a straight line with a negative slope equal to the mass transfer coefficient (K): (13) J = Kin (C.35 shows increasing slopes with increasing recirculation rate (the volumetric flow rate of retentate recirculated back as feed to the channel. it is difficult to carry out a concentration at constant cross-flow velocity (recirculation rate). log concentration curves. Experimentally. The values obtained from Figure 3. According to Equation 13. while the exit region may not be.. This provides a way of experimentally determining the gel concentration.34 show the semilogarithmic variation of flux with concentration for two proteins and two colloidal suspensions.
/
Effect of Concentration. (Usually. ULTRAFILTRATE MEMBRANE BOUNDARY LAYER FEED IN _+ RETENTATE OUT BOUNDARY LAYER
ULTRAFILTRATE Figure 3. because the boundary layer is not well developed at this point. The data of Figure 3.

if the data are not gel-polarized. All data on Figure 3. eventually. the backdiffusive mass transport removes protein from the membrane surface at a high rate.0
I
I
I111111
2
4
6
8 10.36 could be gel-polarized if the average transmembrane pressure were raised sufficiently.Ultrafiltration
173
process stream is concentrated.5 GPM) have lower back-diffusive transport and are gel-polarized.36.36:
Variation from semi-log dependence in non gel-polarized regime.35 and 3. and the flux drops to lower operating lines (see Figure 3.
l. Low recirculation rates K2.0
20 (wt %)
40
60 80 100
Protein concentration Figure 3. In Figure 3. at low pressures or low concentrations. the recirculation rate decreases. it would be expected that operation in this pH range would result in the lowest gel concentration. However. the flux will reach an asymptotic value equal to the pure solvent flux of the membrane.) is less than C. Therefore.
Effect of pH. it will be seen that the gel concentration is 28% in one case and 45% in the other. protein concentrations below 1% at high recirculation rates are not fully gel-polarized. the inconstancy of recirculation rate can lead to inaccurate estimates of the gel concentration (Cc). and not constant. at extremely low solute concentrations. When the data are extrapolated to zero flux. some curvature is expected. both showing flux data on human albumin solutions. In addition.36.35). since changes in pH are not expected to change the
. there is a ceiling on solvent flux due to the limiting hydraulic resistance of the membrane. The concentration at the surface of the membrane (C. Equation 13 will not plot as a straight line on semilog paper. Since the isoelectric point of most proteins is between a pH of 4 and 5. This reflects the difference in solubility of proteins for different conditions such as pH. Further. Recirculation rate
I 11111
1. Comparing Figures 3.

for a given concentration. we can modify the gel concentration (Cc) by altering the pH and/or ionic strength of the medium. suitably modified for mass transfer.
Evaluation of the Mass-Transfer Coefficient It is clear from Equation 10 that the UF flux is determined largely by the mass-transfer coefficient (K).174
Handbook of Industrial Membrane Technology
mass transfer coefficient (K). that option is not permissable.36 do not change in slope but are simply shifted to the left or right depending on the gel concentration (Cc). With proteins and other polyelectrolytes. The mass transfer-heat transfer analogies well known in the chemical engineering literature make possible an evaluation of the mass transfer coefficient (K) and provide insight into how membrane geometry and fluid-flow conditions can be specified to optimize flux.
l3-
12 -
I1 -
g 5 EL ii c
lo-
9-
9t
7 -
Minimum associated with isoelectric paint
3
62
I
4
1
5
6
7
1
9
.
‘PH Figure 3.37: Effect of pH on flux (Cheshire cheese whey). may be used to evaluate the mass transfer coefficient where the laminar parabolic velocity profile is assumed to be established at the channel entrance but where the concentration profile is under development down the full length of the channel. the optimization of flux is largely effected through the parameters that effect the mass transfer coefficient. the operating lines in Figures 3. Thus. The Graetz or L&que solutions25r26 for convective heat transfer in laminar flow channels. Figure 3. Notice the minimum at the isoelectric point. For all thin-channel lengths of practical interest. for most process streams.35 and 3. there can be dramatic change in flux with pH. Levgque’s solution26 gives:
.37 shows the variation in the UF flux from cheese whey with pH. this solution is valid. However. Therefore.% Laminar Flow.

any fluid management technique which increases the fluid shear rate (i) at the membrane surface will increase the flux. This has been confirmed for a large number of solutions ultrafiltered in a variety of channel geometries (see Figure 3.23 Re0. for a fixed bulk stream concentration (Cu). the channel length (L) is important because the boundary layer develops as fluid moves down the channel. In practice. Turbulent Flow. the flux should vary directly as the cube root of the wall shear rate per unit channel length.176
Handbook of Industrial Membrane Technology
cient) may be increased by increasing the channel velocity (U or Q) or by decreasing the channel height (b).1 .38). Consequently.1. 0. In more general terms.
/
.8 SCO.* * . Short channel lengths mean a high cost per unit of membrane area. Perhaps the best known heat-transfer correlation for fully developed turbulent flow is that owing to Dittus and 8oelter.08 0. 0.27 The mass transfer analogy based on the Dittus-Boelter correlation is:
(19)
Sh = 0. 140
I
“6 &
-7
= s 3
.~~
. this is not cost effective since a significant portion of the cost of manufacturing a module is associated with potting or connecting the channels/tubes to a common “header”.1 0.
I
In laminar flow.2. Indeed. .38: Experimental confirmation of 0.06 -
I
.7
I
100 1000 WALL SHEAR RATE 6RARREL LENGTH 1 f/L9 (set-cm)-’ Figure 3. *. Equation 18 shows that in laminar flow. This would argue for shorter channel lengths. the average flux is lower.33 power dependence of flux on wall shear rate/channel length (in laminar flow).2. Longer channels have a thicker boundary layer (at the end) which reduces the back diffusive transport.2: .4 .0.3 _ c E 0.

8 power and inversely with channel height. Measurements of fluid velocity versus pressure drop show a definite transition from laminar to fully developed turbulent flow at Re = 2000.8 power. whereas the 30-mil channel data were taken in turbulent flow. the entrance geometry is such that fully developed turbulent flow occursat much lower Reynold’s numbers. however.OOO. the Dean number (De) governs the transport processes in coiled tubes and channels:
De = Re J + where a and R are tube and coil radii respectively. Dravid et al’e has shown experimentally that the heat-transfer coefficient in a coiled tube varies as De0’55 in
. Again as in laminar flow. It then recirculates inward along the walls of the channel (see Figure 3.e. For flat rectangular channels where dh = 2b. the flux (or mass transfer coefficient) may be increased by increasing the channel flow rate (Q) and by decreasing the channel height (b). most of the ultrafiltration data taken on solutions in turbulent flow are in good agreement with theory. with the Dittus-Beelter correlation. It is well known that coiled tube heat exchangers possesses superior heat transfer characteristics because of secondary flow effects. Equation 20 becomes:
(21)
or
(22)
K = 0. At small curvatures. centrifugal forces tend to throw fluid outward from the center of the channel. Furthermore. The data were taken in a spiral flow thin channel device. The effect is.52 than predicted by theory (0.40)..75 power dependence of flux on recirculation rate (Q) in good agreement with the predicted 0. much more dramatic in turbulent flow-the flux varying with flow rate to the 0. However.Ultrafiltration
177
(20)
K = 0. the turbulent flow data of Figure 3. (The IO-mil channel data were taken in laminar flow.33)-probably because of “secondary flow effects”.) The laminar flow data of Figure 3. Whenever fluid passes through a curved tube or channel.023
It can be argued that any turbulent flow correlation should not be applied for Re <lO. For example.39 have a higher slope (0. the flux is independent of channel length since both the velocity and concentration profiles are established rapidly in the entrance region of the channel. i.02
As in laminar flow. in current thin-channel ultrafiltration devices.39 show a 0.

the experimental values were still well within 25% of the theoretical values. Figure 3.OL 1
I
2 3 4 5 6 7810
20
30
L .52 compared with the theoretical slope of 0. The more Crucial test of the theory.Ultrafiltration Theoretical Prediction of Flux
179
The success of the L$que and Dittus-Boelter relationships in indicating the variation (power dependence) of ultrafiltrate flux with channel geometry and fluid velocity for macromolecular solutions is gratifying.36.) was determined experimentally from data like those in Figure 3. ULTRAFILTRATION ALWilN OF WAN LTC-1 WI%)(15 MIL CHANNEL)
TRANSMEWBR!NE PGESSUBE
30 P-SIG AVG..41: Theoretical vs experimental values of flux with recirculation rate (in laminar flow). spiral flow channel plates). Figure 3. Macromolecular Solutions.41 presents laminar flow data from 15mil channel tubes compared with theoretical values using a diffusivity of 6 x lo-‘cm’/sec and a gel concentration of 45% as determined from Figure 3.
4050
Figure 3.e. .39 shows that agreement between theoretical and experimental values is good in at least some cases (i. is whether these relationships can be used to calculate quantitatively the Ultrafiltrate flux knowing the channel geometry. The accuracy of the L&&que and Dittus-Boelter relationships has also been verified in linear thin-channel tubular equipment ultrafiltering protein solutions. Again.29 The gel concentration (C. The diffusivity of albumin (D = 6 x IO7 cm’/sec) was obtained from a handbook. . of considerable interest to the design engineer.
DROP
GECIRCUIATION
GATE (GPtl)
l.35. Although the slope in laminar flow was 0.33. experimental and theoretical values agree within 25% although the theoretical values are consistently lower. Agreement in the turbulent regime is more striking.. Equation 17 for laminar flow and Equation 22 for turbulent flow were used to calculate the mass-transfer coefficient (K).
.a. fluid velocities and solute characteristics.

experimental flux values are often one to two orders of magnitude higher than those indicated by the L&e^que and Dittus-Boelter relationships. the dependence of flux on diffusivity to the 0.5%.
Other macromolecules in solution seem to fit the gel-model as well as protein solutions. 20 PSI = T.65 PSI Poll.43. The larger.87 power is confirmed..7 1. The theoretical curves are 15 to 20% below the experimental data in both cases. respectively. . The agreement between theoretical and experimental ultrafiltration rates for macromolecular solutions can be said to be within 15 to 30%. The theoretical values were calculated using a diffusivity of 5 x lOI cm2/sec3’ and a gel concentration of 7. For colloidal suspensions.42. Here.23'C 1
40
30
20
10
PROTEIN CONCENTRATION X WGT. 6 7 8 910
I
20
30
40 !a
Figure 3.42:
Variation in flux with diffusivity of retained solute. data from the literature3’ on the ultrafiltration of polyethylene glycol (Carbowax 20 MI solutions in one inch diameter tubular membranes operating in turbulent flow are plotted along with theoretical flux values. For example. Thus.0 2 3 I 45 I III.). Colloidal Suspensions. I 01.180
Handbook of Industrial Membrane Technology
The ability of the L&&que solution to adequately describe the variation in flux with diffusivity of the retained solute is illustrated in Figure 3. albumin data were compared with whole serum data. in Figure 3.
.30 The agreement between theoretical and experimental values is within 14 to 27% at the higher and lower Reynold’s numbers. globulins in whole serum have a lower diffusivity D = 4 x IO1 cm2/secz9 and a lower solubility limit (C.
5-u
CONCENTRATION SERUMPROTEINS OF IN LTClS(15HlL CHANNEL)(PbfiO MHBRANE) RECIRCULATION RATE 23 GPH P.

That such is not the case is illustrated in Figure 3. and this can be seen in Figures 3.. Referring to Equation 24.1111
4 6
I
a I02
2
I Illll
4 6
I I III
a IO3
2 4 6 6
IO’
I04
AVERAGE SHEAR RATE AT MEMBRANE Figure 3. including the 8 /.
z
9
WHOLE BLOOD AND PLASMA UM-IO MEMBRANE.34 and 3.. protein diffusivities of interest in this chapter are as follows: Protein Albumin y globulin Collagen (gelatin) M. should be considerably less than that from the plasma above.34 have a much higher slope (mass transfer coefficient) than would be expected from the latex particle diffusivity. al _
II
I I III
I
I
II. 65..000 170.I7 mils
A WHOLE PLASMA-S I
2
I
I. AP= 10 psi ESMOND CELL.
Likewise..44.000 345.7 x 10-7 x 10-7 x 10 -7
These proteins should exhibit decreasing flux rates in the order given.42. 6 4 0. L=l3 cm ___--------------
WATER FLUX
“E -5
0 WHOLE PLASMA-35 0 WHOLE BLOOD-35 o WHOLE PLASMA-17
mils mils mils mils
n WHOLE BLOOD. one would expect that the ultrafiltration rate from whole blood. Monodisperse polystyrene latexes have both suspension viscosities and
. the polymer latex data of Figure 3.182
Handbook of industrial Membrane Technology
Macromolecules of higher molecular weight will have larger particle dimensions and lower diffusivities.W.r red cells.000 D 20°C cm2/sec.44:
SURFACE
(set”)
UF of whole blood and plasma. For example.

. The discrepancy between the experimental flux value and theoretical laminar flow value is a factor of 38 at 40 GPM and a factor of 15 at 5 GPM! If the assumption is made that these data were taken in turbulent flow (completely unwarranted). Using the values cited above for the diffusivity (D) and the gel concentration (C. This is in good agreement with other values reported in the literature. the theoretical flux is plotted for the 40 GPM recirculation rate in Figure 3.5 at all recirculation rates.TION OF STYRENE BUTADIENE POLYMER LATEX IN LTC-I 15 NIL CHANNELS XEM WIBRANE GO PSIG MR&%_TIW'E"G~NE
\
1009Q8070 GO504030 20 1”
\
\ \
rn-
CONCENTRATION WOL XI 2 3 4 5
THEORETICAL
40 GPH
1
678910
20
30
40 50
Figure 3.45: UF of styrene butadiene polymer latex showing disparity between experimental results and theoretical prediction. Most of the data on this plot (below 30 GPM) are taken at Reynold’s numbers below 2000 and therefore assumed to be in laminar flow.
Flux vs.Ultrafiltration
183
sedimentation coefficients in good agreement with the predicted Stokes-Einstein diffusivity. the theoretical flux would still be less than 5/sof the experimental value!
13cnollflr -_ \ \ '\
RECIRCUldTlON RATE 40 GPN
ULTRAFILTRA. the slope is more nearly that indicated by theory. The Stokes-Einstein diffusivity was calculated to be 2.).46.3 x 10-s cm*/sec for a carboxylic modified styrene-butadiene copolymer latex which had an average particle size of 0. The gel concentration of this material has been determined to be 75% for numerous thin-channel runs in both tubular (see Figure 3. recirculation rate data for a thin-channel tubular unit are plotted in Figure 3. but the discrepancy is still a factor of 7.45) and spiral flow equipment.19 p.45 (thin-channel tubes) and found to be l/38 of the experimental thinchannel tube value! Even if the calculated diffusivity were an order of magnitude larger. Theoretical values are also plotted assuming laminar and turbulent flow.

However. The experimental slopes tend to be higher than the theoretical slopes in both laminar and turbulent flow except for the laminar flow data at 1% concentration which follows the predicted ‘h power dependence.47).194
Handbook of Industrial Membrane Technology
101 l81 61 41
ULTNAFILTNATION STYRENEBUTADIENE OF POLYMER LATEX IN LTC-1 (15 NIL CHANNEL) XMSO
21
11
I f
1
1
1
1
I
/2 3 4
RECIRCULATION RATE (GPM) 5 6 78910 20 30 40 70
Figure 3.
The gross discrepancy between theoretical and experimental flux values also exists in data obtained from spiral flow thin-channel equipment (see Figure 3. suggesting that the data labeled “turbulent flow” may be in a more nebulous transition region between laminar and turbulent flow. The experimental laminar flow values are a factor of 19-29 higher than the theoretical values whereas the experimental turbulent flow values are a factor of 8-15 higher than the theoretical values. Again. as calculated from theory.
. there is no abrupt change in flux. In these data.46: Flux vs recirculation rate in linear thin channels (styrene butadiene polymer latex). the gross failure of the theory in estimating experimental values is evident. the transition from laminar flow is clearly seen (change in slope) at Reynold’s numbers only slightly below 2000.

rate in spiral thin channels (styrene butadiene
An evaluation of more than 40 different colloidal suspensions in our laboratories has indicated that the diffusion coefficient calculated from the ultrafiltrate flux using the L&v$qque Dittus-Boelter relationships is generally from one or to three orders of magnitude higher than the theoretical Stokes-Einstein diffusivity.47: Flux vs recirculation polymer latex). TRANSMEHBRANE PRESSURE DROP
100
200
400
600
1000
2000
4000
Figure 3.Ultrafiltration
I
I I
185
I I
Illll
ZOC-
ULTRAFILTRATION OF STYRENE BUTADIENE LATEX TCl-D (30 NIL CHANNEL) (PM30 WIBRANE) 40 PSI AVG. It is evident from the above that minor adjustments in molecular parameters
.

188
Handbook of Industrial Membrane Technology
such as diffusivity. For example. with the assumption that the proper diffusion coefficients are the Stokes-Einstein diffusivities for the primary particles. kinematic viscosity. He noted that the region immediately adjacent to the walls of blood capillariestends to be free of blood cells. the layer will continue to grow until the channel is completely full of 75% solid material-resulting in a drop in recirculation rate with time. Actual ultrafiltration fluxes are far higher than would be predicted by the mass transfer coefficients estimated by conventional equations. Similar discrepancies were noted by Blatt et a13’ for colloidal suspensions such as skimmed milk.
. Blatt concluded that either (a) the “back diffusion flux” is substantially augmented over that expected to occur by Brownian motion or (b) the transmembrane flux is not limited by the hydraulic resistance of the polarized layer. this is called the “tubular pinch effect”. The apparent viscosities of such suspensions vary with tube radius. or gel COnCentratiOn are incapable Of rasolving order of magnitude discrepancies. and the flux should be proportional to the transmembrane pressure drop. polymer latexes. Palme? was able to skim off a plasma rich layer at thewall through fine branches and was able to measure increases in hematocrit from near the wall to the axial region. However. and flow rate. and clay suspensions. the membrane must be. length. Thin channel tubes and spiral flow modules running continuously at constant latex feed concentration and pressure drop for periods approaching one year have shown no decreases in recirculation rate or accumulation of polymer latex in the channels. He favored the latter possibility. arguing that closely packed cakes of colloidal particles have quite high permeabilities. Tubular Pinch Effect.
(2)
These observations lead to the conclusion that the back-diffusive transport of colloidal particles away from the membrane surface into the bulk stream is substantially augmented over that predicted by the L&v$que or Dittus-Boelter relationships. Experimental data deny this-showing threshold pressure (above which flux is independent of pressure) (see Fig ure 3.48). If the gel layer is not the limiting resistance to flow. this is not a plausible hypothesis for the following reasons: (1) If the gel layer is not the limiting resistance to flow. It is known that colloidal particles flowing down a tube tend to migrate across the velocity gradient toward the region of maximum velocity. The lateral movement of particles across the streamlines in laminar flow was first observed and recorded in 1836 by Poiseville. blood flowing through fine glass capillaries reaches an equilibrium state in which the red cell concentration in the tube is less than that in the inflowing or outflowing blood-presumably the result of axial drift of red cells and their consequent faster average transit than plasma. casein. This phenomenon also explains why many colloidal suspensions exhibit less frictional pressure drop than would be expected from the fluid viscosity.33 To account for such anomalies it has been postulated that there exists a lubricating particle-depleted (“plasmatic”) layer at the wall of vessels in which there is a nonuniform shear field.

The “tubular pinch effect” has also been observed experimentally for suspensions of rubber disks. elastomer filaments. At this position. carbon black.11. PVA spheres. polystyrene spheres.48:
Independence of flux on pressure (styrene butadiene polymer latex). and Silberberg35~36-working with dilute suspensions of rigid spheres.
.
Segr.) to the total number of particles (C total) in the entire channel width. insoluble salts. aluminum particles. The observations of Segr6 and Silberberg have spawned a number of theoretical and experimental studies investigating the effect and analyzing the cause. glycerol and silicone oil in various continuous flowing media. where the particles migrated away both from the tube wall and the tube axis reaching equilibrium at an eccentric radial position. were the first to publish their observations of the “tubular pinch effect”. One of the more spectacular visual studies was made by Brandt and Bugiarello.8 GPH
1X LATEXCONCENTRATION
AVERAGETRANSREH8RANE PRESSURE DROP (PSI) 10 20 30 40 9 GO 70 80 90
Figure 3. the spheres became regularly spaced in chains extending parallel to the tube axis.37 They made direct photographic observations of small Dylite spherical beads suspended at concentrations of 1.Ultrafiltration
187
11
90
I
VARIATION FLUX WITH TRANSMMMNE OF
PRESSURE DROP
STYRENEBIJTADIENE POLYt'!ER LATEX IN LTC-1 (1s UlL CHANNEL~MiM IIEMBRANEI ATREIATIVELY RECIRCULATION LOW RATE .7 to 5% in a glycerin-water solution and flowing in laminar flow (Re from 400 to 1640) through long rectangular channels made of Plexiglass. aluminum-coated nylon rods.49 gives the average half-channel particle distribution (obtained photographically) expressed as the ratio of the number of particles per band (C. Figure 3. The ratio of the channel width to the bead diameter was 25.6.

5 1). The tubular pinch effect can explain much of the anamolous UF data for colloidal suspensions. Brandt and Bugliarello found that the process was accelerated by increases in flow rate and delayed by increases in average concentration. will experience a pressure difference across its diameter due to the higher velocity on one side than on the other. Equations 25 and 26 both predict a radial migration velocity (VI increasing as the square of the cross-flow velocity (U) [The Reynolds Number (R.g.75 in turbulent flow (e. the water flux through the porous wall will still carry particles to the wall.50 and 3. With UF.33 in laminar flow and 0. Colloidal particles are large enough that any given particle.
.Ultrafiltration
189
The plots show very clearly the migration of particles to the center of the channel as the distance (Y) from the inlet of the channel increases. the particle will experience a “lift-force” tending to move it away from the wall toward the center of the channel..They used Stoke’s law for the neutrally buoyant case:
(25)
where
V is the radial migration velocity U is the average fluid velocity down the channel or tube Re is the Reynold’s number t-o is the particle radius R is the tube radius r is the radial position of the particle in the tube F(r/R) is a function of the radial position of the particle in the tube or channel. The Navier-Stokes equation has been solved By Cox and Brenner (unpublished data) who computed the lateral force required to maintain a sphere at a fixed radial position (t-1. see Figures 3.) also includes U). This can explain why the dependence of UF flux on cross-flow velocity (U) can be higher than 0. but the “lift” of particles away from the wall (due to the tubular pinch effect) will certainly augment the back diffusive mass transfer described by the L&e^que and Dittus-Boelter relationships.
Equation 25 has the same form as the empirical equation used by Segre and Silberberg% to correlate their data:
where r* is the equilibrium radial position of the particle which was found to decrease as (m/R) increased. Thus. This “lift-force” is not unlike the aerodynamic lift on an airfoil due to the “Bernoulli effect”. not in the center of the tube.

52: Enhancement of protein flux with particle addition. 0
5
10
15
20
25
30
40
PARTICLE LOADING CVOL.. Green and Belfort3’ have combined the equations for particle migration due to the tubular pinch effect with the normal back-diffusive transport to calculate
. He attributed this to (1) the mixing action of the particles and (2) the mechanical scouring of the membrane surface.%) Figure 3. This may explain the larger discrepancies between experimental and theoretical flux values in 15 mil channels (see Figure 3.
Equations 25 and 26 also predict that the radial migration velocity (V) will increase as the tube radius (R) decreases. Thin channels are more effective in depolarizing the membrane surface via the tubular pinch effect. Bixler et a13* found that adding glass and plastic beads (ranging in size from 30 to 100 /J) to a protein solution augmented the UF flux (see Figure 3. the flux with the red cells present is higher than that with plasma alone. The migration of the larger red cells away from the membrane surface tends to drag the plasma proteins along.Ultrafiltration
191
The data of Figure 3. the tubular pinch effect tends to depolarize the membrane surface of red cells yielding a flux similar to that obtained with plasma alone. Presumably.46) than in 30 mil channels (see Figure 3. In some cases.52).44 show similar flux values for whole blood and plasma.47).
I+
I
I
I
I
I
I
6-
8.

For particles bearing an electric charge.85%
0
Ia0
200
380 Centrifugal
400 Force (x g)
580
600
Figure 3. Centrifugal Force. these complex theoretical calculations still come short of accurately predicting the experimental flux from various colloidal suspensions.54 and 3.64% I . thereby augmenting the mass transof fer coefficient (K).55). Many particles and/or macromolecules bear a charge (USUally negative). centrifugal force has been explored.192
Handbook
of Industrial
Membrane
Technology
flux values which are closer to the experimental results. an electric field will always give higher filtration rates than cross-flow filtration alone (see Figures 3.6BX
3. Figure 3. Unfortunately. Augmented Cross-Flow Effects
Other methods have been sought to augment the depolarization of the membrane by cross-flow.eex
7. an electric field has been investigated.53 shows the increase in the mass transfer coefficient (K).81X I . The family of curves are for different concentrations of casein. When the apparatus was spun. centrifugal field strengths from 100 to 600~ resulted in flux improvement factors of 3 to 16. Robertson et aVk’ designed an apparatus in which the centrifugal force vector was perpendicular to the membrane surface but opposite (and parallel) to the flux vector.400 daltons) were passed over the membrane in laminar flow. Solutions of casein and dextran (60.
in mass transfer
coefficient
(K)
with
the imposition
of
Electric Field. Consequently. due to the centrifugal force. the application of an electric field can cause these species to migrate away from the membrane.
.
8. A step-wise iterative procedure was used to calculate a detailed particle trajectory analysis. Since both are negatively charged in aqueous suspensions. as calculated from Equation 10.53: increase centrifugal force. The increase in UF flux depends on the electric-field strength as well as the cross-flow velocity. For particles with a higher density than water. Henry et a141 has investigated cross-flow electro-filtration kaolin clay suspensions and oil-water emulsions.

by the application of voltages above E.5 /. the data of Figure 3. but above the critical voltage.e. breakthrough can be avoided.55). Without the electric field.56 is for E less than E. the voltage becomes critical as we decrease the cross-flow velocity).r pore size membrane. For example. Solutions/suspensions which cannot be ultrafiltered economically due to severe fouling problems might be filtered with this technique if the particles and/or macromolecules bear a charge. most oil droplets in emulsions are between 0.) Effect of Temperature It has been found experimentally for a large number of membrane systems (including MF.Ultrafiltration
195
There are three operating regimes in cross-flow electrofiltration depending on whether the field strength (E) is above. or equal to the critical voltage (E. regime (c). At the critical voltage. membrane researchers often refer to the 3% rule (that flux increases 3% per “C) as a rough rule of thumb. The critical voltage is defined as the voltage at which the net particle migration toward the membrane is zero. In this regime. In Figure 3. The flux decreases with decreasing Reynolds number until a point is reached where the convective transport of particles toward the membrane is just equal to the electrophoretic migration away from the membrane (i. its greatest potential appears to be in the elimination of the gel layer altogether. the lower curve for E = 3. but the flux is higher than normal because the back-diffusive transport of particles away from the membrane is augmented by the electrophoretic particle migration. i.5% for every “C rise in temperature. It is even possible that this technique might make the fractionation of gamma globulin from albumin feasible. below. (Normally. Though cross-flow electrofiltration has not yet been commercialized. Thus.54 and 3. Since the viscosity of water decreases by about 2. increasing the tangential velocity decreases the flux.6 /. where the convective transport of particles toward the membrane is exactly equal to the electrophoretic migration of particles away from the membrane.I in diameter. increasing the tangential velocity is expected to have no influence on the flux because fluid shear can only improve the transport of particles down a concentration gradient. Further decreases in cross-flow velocity will not decrease the flux as there is no concentration polarization. there is no concentration gradient.54 were taken using a 0.9 v/cm shows a transition in slope. because the diminished thickness of the boundary layer increases the diffusive transport of particles down the conventration gradient toward the membrane (see Figures 3..55. In this case. the oil droplets would not be retained by the larger pore sizes in MF membranes. In this regime.. there is no concentration polarization because the electrophoretic transport is equal to the convective transport.). the electrophoretic migration of particles away from the membrane is greater than the convective transport to the membrane. Regime (a) in Figure 3.. Oil-water emulsions might be separated by higher flux MF membranes. UF and RO) and feed streams that the permeation rate is inversely proportional to the fluid viscosity.
. When the voltage is greater than the critical voltage.e.1 and 0. increases in the tangential velocity will diminish the boundary layer thickness resulting in a higher flux. the concentration of particles near the membrane surface is depleted due to removal by the electric field. regime (b). Concentration polarization still exists. In this case..

. Figure 3. the lower slope of the IOWtemperature data was due to a lower recirculation rate due to the increased viscosity of the process stream). However. the inverse viscosity rule does not apply for all points on the graph. (Incidentally.?
= pressure drop
= fluid viscosity = pore length (including a tortuosity factor)
The only variable in Equation 27. The mass-transport coefficient (K) and the gel concentration (C. the ratio of the flux at 25°C to that at 4°C goes to infinity. is the viscosity:
where AE. the membrane itself is no longer the limiting resistance to flow when “gel-polarized”. is the activation energy (3750 Cal/g-mol for water). Clearly. Poiseuille’s law predicts that flow (J) through any porous media should be described as follows: (27) where N d Nnd4AP ‘= iz?jYL
= number of pores per square foot = pore diameter (average)
AP
p !. which is temperature dependent. The flux is given by Equation 10 instead of Equation 27. These results are not unexpected for the transport of pure water through porous membranes.
SOLIDS CONTENT
(weight percent)
Figure 3.57:
Effect of temperature on UF of a high-protein meal. as we have seen. At 45% solids.196
Handbook of Industrial Membrane Technology
Pure Water Transport.57 is an illustration of how different protein solubilities at 4°C and 25°C will shift the gel-concentration.) vary with temperature.

9.58: CW6?% 1
Effect of temperature on UF of bovine serum. 21.58). the higher temperatures tend to denature the proteins (reducing their solubility). the effect of temperature on gel concentration is complicated and often unique for the solute/solvent system being filtered. recirculation rate of 150 cdmin at 25 psig
IO -
50 CONCENTRATlOll?OF PROTEIN Figure 3. ll-mil channels.Ultrafiltration
197
On the other hand.
Differentiating Equations 16. some proteins seem to show a gel concentration invariant with temperature (see Figure 3.r(cm2 of membrane. In this case. 24 and 28 yields the following expressions for the dependence of flux on temperature: For laminar flow gel polarization
dJW
(29) dT
For turbulent flow gel polarization
(30)
. Thus. thereby compensating for the expected increase in solubility with temperature.

This results in an overall long-term decay.198
Handbook
of Industrial
Membrane Technology
The derivation of Equations 28 and 29 may be found in Reference 17-Ap pendix E. there is a gradual flux decay with time as in Figure 3. acids.59. In effect. the flux cannot always be restored to the initial value. It assumes that C.
OO
I
I
I
40
I
I
80
I
TIME
I
120 (days)
I
. For most applications. However. When these equations are used to calculate the percent variation of flux with temperature near room temperature. Rather it is the result of the accumulation of materials on the membrane which no longer participate in the mass-transport to or away from the membrane.
160
I
I
200
I
Figure 3. pp. they “blind” small sections of the membrane. Some polymers have a higher susceptibility to fouling and chemical modification of the membrane surface can have a profound effect on the propensity to foul. such as detergents. or even organic solvents. 2-97 to 2-99. even with periodic cleaning. This is not due to internal pore fouling (as in symmetrical MF membranes). The advantage of a chemically-resistant membrane is that severe cleaning agents may be used.
Often preventive measures may be taken to avoid fouling the membrane. is invariant with temperature. If fouling does occur. the results are usually within 25% of that predicted by the inverse viscosity rule. thereby reducing the effective area and the flux through the membrane.
. Prefilters or screens can be used to remove large particles which block thin channels or accumulate in stagnant areas of the module.
FLUX DECAY AND RESTORATION
With some process streams the flux can be stable for months or even years without cleaning or membrane replacement. Low pressures avoid compaction of gels on the membrane. the membrane deposits can sometimes be removed by aggressive cleaning agents.59:
Long term flux decay and restoration
by cleaning. however.
MEMBRANE
FOULING. bases. High cross-flow velocities tend to sweep deposits away.

Figure 3.Ultrafiltration
Effect of Cross-Flow Velocity
199
High cross-flow velocities tend to prevent fouling and also aid in the cleaning process.
FLUX DECAY
J Flux
10
100
TIME
Figure 3. medium and high recirculation rates (QL.60:
Effect of cross-flow velocity on long term flux decay. 10 days of data often permits extrapolation to 100 or even 1. QM and QH). Plotting the data in this way often permits a reasonably good extrapolation of the flux for much longer timesup to 2 or 3 years. Figure 3.
Figure 3.14 I/s: -A0.
High cross-flow velocities also facilitate cleaning. -a1.
.614’ shows that the flux is restored more rapidly and to a higher level with high velocities.61:
Effect of recirculation rate on detergent cleaning.60 shows the flux decay on a log-log plot for low. and because of the cyclical nature of the log-log plot.57 I/s) and time on flux during cleaning of tubular ultrafiltration membranes.000 days.
I
o-25
Cleanmg time (h) The effect of detergent circulation rate (-O1. The flux decay data usually plots on a straight line.51 l/s.

and sputtering. butyl Cellosolve). For example. the inorganic coatings tend to resist fouling producing a more stable flux.62 may show greater flux stability at pressures (PL) lower than the threshold pressure. the XM-50 membrane experienced catastrophic flux decline-presumably because of the electrostatic interaction with the positively charged paint particles. the long-term flux at this pressure may be higher. However. When the electropaint users switched to cathodic paints.g. Membranes made from hydrophilic polymers like cellulose acetate are generally less prone to fouling than the hydrophobic polymers. However.
. plasma polymerization. Cellulose acetate membranes provided a stable flux but could not be cleaned with the solvents used in paint makeup (e.200
Handbook of Industrial Membrane Technology
Effect of Pressure If no fouling occurs. In this case. In some cases. for solutes which form semigels on the membrane. Flux decay data like that of Figure 3. organic solvents and elevated temperature.
Effect of Membrane Surface Treatments Changes in cross-flow velocity or transmembrane pressure cannot always alleviate fouling. cellulose acetate is limited in its tolerance to high or low pH. the maximum flux will be obtained in the gel-polarized regime above the threshold pressure (PT). surface modification of the more chemically-resistant polymers has rendered them less susceptible to fouling.64 shows an increasing resistance of the protein gel-layer (I$) (decreasing flux) with time for three samples of the same membrane. a “charged” XM50 membrane (designated CXM) was developed for cathodic paints. Figure 3. pressures (PH) higher than the threshold pressure may compact the gel layer resulting in greater fouling. Therefore.62: Effect of pressure on long term flux decay. FLUX DECAY
J
Flux
10
100
TIME Figure 3.63 compares the performance of the “charged” membranes with the uncharged. a vinyl copolymer membrane (Romicon’s XM-50) was used for years with anodic paint.. Various functional groups may be applied to a membrane by techniques such as chemical grafting. two which have been coated with carbon and a polysiloxane. Even though the initial flux at PL is lower (since not gel polarized). The slight electronegativity of this membrane repelled the negatively charged anodic paint producing stable flux over a long period of time. Figure 3.

65 shows the effect of papain and fungal proteinase in decreasing the rate of flux decline over a 20. Figure 3.0%. The enzyme may be crosslinked and immobilized on the membrane surface with 0. (m) fungal proteinwe P. (-1 model. There are currently four generic configurations for UF membranes in industrial use: tubes. Modelling flux of prototype and control/effect of different enzymes: BSA. (A) Control. Equations 17 and 22 show clearly that thin channels promote higher mass transport and flux.5% albumin or hemoglobin.
c
PM-10 membrane.202
Handbook of Industrial Membrane Technology
This resulted in a 25 to 75% improvement in the permeate yield during a 22-hour run concentrating 0. In addition.125% glutaraldehyde in a phosphate buffer (pH 6. hollow fibers. Tubes Perhaps the simplest configuration is a tube with the membrane cast on the
. Hl. but this must be balanced against their greater propensity to foul. 11. ease of cleaning and replacement. Effect of immobilized protease on UF of BSA (a self-cleaning
MEMBRANE
CONFIGURATION
The design of the membrane package or module and the fluid management within that module will profoundly affect membrane performance.65: membrane). H2. the optimum design for one application may be totally unsatisfactory for other applications.5 to 5. (A) papain (Corolase SlOO). high performance modules must be evaluated in terms of cost.5hour run concentrating BSA from 0. For example. plate and frame units. Figure 3. The net protein loss through the membrane due to cleavange of albumin by the active enzyme was found to be only 4% of the total. and spiral wound modules. Further.5).

Tubes may be effectively cleaned by introducing a number of sponge rubber balls slightly larger than the tube diameter (see Figure 3.68 is a schematic diagram of a device to dispense and recover these balls automatically at preset intervals. Periodic passage of these balls can prevent severe flux decline (see Figure 3. Membrane
supporter
Membrane
Sponge ball
Figure 3.67:
Effect
of sponge ball cleaning
. When the membrane is spent.66) into the process stream. Figure 3. usually the whole tube must be replaced. The porous support tube (often fiberglass reinforced epoxy) is the dominant cost factor and the membrane area per foot of tube is low.6). a cleaning solution may be used in conjunction with the sponge balls.
-
0 With Sponge
l Without
Ball Cleaning Ball Cleaning 8 lo 12 Time 14 (hr) on long term flux decay.
Removed
deposit
. Tube diameters from 0. cleaning can be done simultaneously with processing.Ultrafiltration
203
inside wall of the tube (see Figure 3. 16 18 20
22
Sponge 4 6
1’
2
Operating
Figure 3. When severe fouling is encountered.66:
Schematic
of sponge ball cleaning. The biggest disadvantage of tubes is their cost.67).25 to 1 inch are less prone to foul and more easily cleaned than any other configuration. Others have developed epoxy bonded consolidated sand as the porous support in an attempt to reduce costs.
1 -
I
. In this way. PCI utilizes a “paper” insert on which the membrane is cast as the replaceable element.

# CONCENTRATE (RETENTATE)
BRAIDED
SUPPORT
ULTRAFILTRATE
FEED STREAM
ULTRAFILTRATION MEMBRANE
Figure 3. The tube is braided on the outside for strength.te
Pressure
Concentrate
Figure 3.204
Handbookof Industrial
MembraneTechnology
-Ball
positioning handle (Available in either manual or automatic mode)
Permea. the splined core in Figure 3. This makes a fairly inexpensive tube.69:
Splined
core linear thin channel.
. but the packing density (membrane area per unit volume) is still quite high and there is considerable “parasite drag” due to the splined core.
The trend is toward using smaller diameter tubes or volume displacement rods to reduce the volume of fluid pumped per unit area of membrane. For example.69 has sheet stock wrapped around the core and sealed longitudinally.68: Sponge ball dispenser
Feed
and collector.

D. shorter modules have a higher cost per unit of membrane area. The lower burst-pressures associated with these large diameter fibers generally limit the maximum diameter to below 2 mm I. The same technique may be used in the field to reclaim leaking modules. Unfortunately. etc. the back-wash fluid should be filtered to remove any particles which would tend to lodge in the porous wall of the fiber. However. low manufacturing yields have kept selling prices ($/sq ft) equal to or greater than that for spiral-wound modules. There is no “parasite drag” and no expensive porous support tube.Ultrafiltration Hollow Fibers
205
Conceptually. Because hollow fibers are self-supporting and hold up well under the compression force of a reverse transmembrane pressure drop.
Potentially. plate and frame units. A gas line is used to pressurize the permeate port on the shell of the module with just a few psig. Fortunately.5 mm I. There are several methods under development for increasing the burst-strength of hollow fibers to permit higher operating pressures. These fibers are plugged using hot-melt. The open-ended module is immersed in a tank of water. they can easily withstand back-wash pressures of 15 to 20 psi. and spiral wound modules do not recommend back-washing due to problems with membrane delamination and glue line seal rupture. since most of the manufacturing costs are associated with potting the fibers at the ends of the module (see Figure 3. hollow fibers are the ideal membrane configuration. The biggest disadvantage is the pressure constraint which limits the crossflow velocity down the lumen of the fiber. The end of the module is observed to pin-point bubbling fibers (leakers).
Filtrate
hollow-fiber
membrane
Hollow
fiber
terminal 1140mm
Figure 3.D. the hollow fiber configuration is more susceptible to fouling and plugging than any of the other three configurations.70:
Schematic of hollow fiber module. Reject modules with only a few fiber leaks can be repaired using a “bubble point test” (see Chapter 2).) are becoming popular to improve fouling resistance. hollow fibers may be cleaned by back-washing which tends to compensate for their propensity to foul. hollow fibers should be the most economical membrane configuration available. The module is then turned upside down and the test repeated to plug the leaking fibers on the other end. stainless steel nails with epoxy.
. The fibers may be pressurized on the inside (up to 30 psig) permitting “thin-channel” fluid management of the feed stream (refer to Figure 3. In general. Larger diameter fibers (up to 2. Shorter fiber lengths permit higher velocities at maximum inlet pressures of 30 psig. Manufacturers of tubes.13). However.70).

permeate will flow in the backward direction from the shell side to the inside (lumen) of the fibers. until the pressure builds up to the average of the inlet and exit pressures inside the fibers. Unfortunately.7 1) and more aggressivecleaning solutions must be used. If the modules are installed with auxiliary piping so that the process stream flow direction may be reversed. The reason is that hollow fibers often operate in laminar flow. but in the exit half of the module. This gentle back-wash of the exit half of the module often recovers up to 80% of the flux decay. This is called “cleaning by recycling”.206
Handbook of Industrial Membrane Technology
The permeate itself is often the ideal back-wash fluid since it has been through the membrane once. permeate by itself will not always restore the flux (see Figure 3. The permeate will collect in the shell side casing of the module. the sweeping action of the process stream along the surface of the membrane effectively removes material loosened by the gentle back-wash. An auxiliary pump and filter must be provided to deliver the cleaning solution under pressure through the permeate port to the shell side of the module.
Water flux after backflushing with
NaOH
Initial water flux Water flux
after bachwashing with water
Transmembrane
pressure difference.
lb/in’
Figure 3. Cleaning study showing the effect of back-flushing with sodium hydroxide on recovering the flux of a cartridge intentionally fouled with cultured skim milk. The materials back-flushed will exit from the fiber lumen and must bedischarged from the system so as not to redeposit on the walls of the fiber.72). The inlet half of the module permeate will flow in the normal forward direction through the membrane to the shell side.
. Further.
Perhaps the simplest cleaning technique of all is to continue circulating the process stream through the module with the permeate port(s) closed (see Figure 3.71: Restoration of flux in hollow fibers by back washing. and the gel layer is most fully developed in the exit half of the module. the other half of the module may be back-washed.

77.208
Handbook of Industrial Membrane Technology
Plate and Frame Units Flat sheet membranes in a plate and frame unit offer the greatest versatility of any configuration but at the highest capital cost. There are many variations on the plate and frame theme. If a leak develops in one membrane. In many of these units.73) and Rhone-Poulenc (Figure 3.
.73 to 3. One of the advantages of the DDS (Figure 3. Alternatively.73:
DDS plate and frame design. virtually any membrane may be cut to the appropriate shape and installed in the unit. but the replacement labor is high. the permeate tube from the leaking membrane may be pinched off since the loss of one plate out of 50 to 100 plates is not a significant percentage of the flow.
Figure 3.74) designs is that each support plate has a separate permeate port. Representative designs are shown in Figures 3. Disassembly also facilitates cleaning. The membrane replacement costs are potentially low. it can often be located visually without disassembling the whole stack.

Millipore Corp.47 and 2.76) heat seals the membrane to plates which drain off filtrate through a central permeate port to the outside of a pressure housing. the channel plates are preferred over the
. Alternating holes on the edge of the cassette carry retentate or filtrate to the appropriate manifold. When filtering cellular or colloidal suspensions. and New Brunswick Scientific Co. The "cassette" is a membrane packet (schematically shown in Figure 3. The cassettes are separated by screens or channel plates (see Figures 2. sell a "cassette" plate and frame system.48 in Chapter 2).
The Dorr Oliver design (Figures 3.Ultrafiltration
209
Figure 3.74:
Rhone-Poulenc plate and frame design.75 and 3.77) comprised of two membranes enclosing a filtrate collection screen (sealed around the edges).

Flat sheet membranes with alternating gasket screens and channel plates may be used in lieu of the cassettes only if the lateral permeability through the membrane is low enough to prevent contamination of the filtrate with the retentate or leakage out the edges of the stack. Since the membranes operate in parallel. the cross-flow velocity across each membrane will diminish as the membrane area is increased.210
Handbook of Industrial Membrane Technology
screens which collect particles on the cross members of the screen.
.75:
Dorr-Oliver flat-plate membrane element. The maximum number of membranes in such a stack is limited by the volume of flow which can be supplied through the manifold.76:
Dorr-Oliver assembled flat-plate cartridge.
Figure 3.
Figure 3.

Ultrafiltration
.

A spiral wound module is essentially the cassette of Figure 3.78:
Spiral wound
module
unrolled.D. this eliminates the stagnant area and facilitates in-place cleaning and sterilization. An envelope of two membranes enclosing a filtrate carrier is sealed around three edges and the fourth edge is connected to a perforated tube which carries the permeate (product water) (see Figure 3.
. for sanitary applications. the corrugated spacers are preferable for cellular suspension processing. How_ ever. a screen (controlled bypass spacer) is used to permit flow in the annular region between the module and the pressure vessel. They currently provide one of the least expensive UF modules available in terms of cost per unit of membrane area.79).77 rolled up in a “jelly-roll” configuration. As the module is rolled up.
PRODUCT
WATER
PRODUCT WATER FLOW (AFTER PASSAGE THROUGH MEMBRANE 1
PROOUCT WATER SIDE BACKING MATERIAL WITH MEMBRANE ON EACH SIDE AND GLUED AROUND EDGES AND TO CENTER TUBE. The whole module is inserted in a pressure vessel.D04 MEMBRANE \
MATERIAL MEMBRANE BRINE SIDE SPACER
I
Figure 3. Again. the membrane layers are separated by a screen or corrugated spacer where the feed solution flows parallel to the tube axis (see Figure 3.212
Handbook
of Industrial
Membrane
Technology
Spiral Wound Modules Spiral wound modules were originally developed for RO but are capturing an ever increasing share of the OF market.78). sometimes several modules are placed in one long pressure vessel. Often a chevron seal is used to seal the outer surface of the module to the inside of the pressure vessel. this insures that all of the feed stream is forced between the membrane layers.
.

Table 3..80 shows the UF modules operating in parallel. microorganisms. Although the single pass operation in Figure 3.g. and organic molecules in the concentrate can be sent to the drain as “blow-down”. They are more prone to fouling than tubes and some plate and frame units (depending on the type of feed channel spacer).80.3: Tubular UF Membrane Configurations
Hollow Fiber Low Moderate
Plate & Frame High Low
Spiral Wound Low Moderate/ Low Good
Cost/Area Membrane Replacement cost (not includinglabor) Flux (GSFD)
High High
Good
Fair/Poor
Excellent/ Good Good/Fair
Packing Density (Ft2/ft3) Hold-up Volume
Poor
Excellent
Good
High High
Low Low
Medium Medium
Medium Medium
Energy Consumption Fouling In Place
Excellent Excellent
Poor Good
Good/Fair Fair/Poor
Good/Fair Fair/Poor
UF PLANT DESIGN Mode of Operation The arrangement of the membrane modules and their mode of operation can affect the economics as much as the module design. it is sometimes advantageous to operate in series to maintain a longer contact time with the membranes and to reduce the volumetric
. deionized water). the colloidal silica. Single-pass UF can only be used on relatively pure water streams (e. and most cannot be autoclaved. These are shown in Figure 3.3 compares the relative advantages and disadvantages of the four basic configurations. In the case of the hundred-fold concentration of a dilute hormone. In the case of deionized water. Table 3. 99% of the feed will pass through the membrane.214
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of Industrial
Membrane Technology
Spiral-wound modules cannot be unwrapped for cleaning lest the glue line seal rupture. If the retained species are low in concentration. There are basically 3 different operating modes for UF. the “recovery” of permeate can be as much as 95 to 99% of the feed. but they are more resistant to fouling than hollow fibers.

*
RMEATE
A. BATCH ULTRAFILTRATION
(. Recirculation of the process stream across the membrane is necessary to obtain the desired concentration or recovery.Ultrafiltration
215
pumping requirement.LJr
l BLEED
FEED *
B. Typ ically the ratio of recirculation rate to permeate flux is 10 to IOO-fold.
. the flux is too low to operate in the single-pass mode. the recovery of permeate in a single pass is a small percentage of the feed (very little concentration of retained species). SINGLE-PASS Figure 3. The plate and frame units can often be arranged to operate with the membranes in parallel or in series. FEED AND BLEED ULTRAFILTRATION
C. In most cases.80:
ULTRAFILTRATION UF operating modes. Recirculation can be accomplished with either a batch concentration or with a feed and bleed operating mode.

Indeed. the feed stream is fed into a recirculating loop. the volume in the reservoir goes down and the concentration of retained species goes up. salts) remain at the same concentration in the reservoir and in the permeate stream. a higher average flux is achieved due to the gradually increasing concentration. Optimum Recirculation Rate There is an optimum recirculation rate for each stage. In a multistage feed and bleed system. it is advantageous to operate as many stages as possible at the lower concentrations. (See Chapter 6: Process Design and Optimization. The total membrane area required in a feed and bleed mode can be reduced considerably by arranging the modules in two or more feed and bleed stages. In this multistage (cascade) system. but the utilities costs go up. equal numbers of modules for each stage often comes close to the optimum configuration. the third stage from 4 to 7%. Freely permeable species (e. Often a ratio controller is used to keep the feed to bleed ratio constantequal to the volumetric concentration ratio required. the feed and bleed mode is less efficient due to increase in entropy upon mixing the low concentration feed with the high concentration recirculating retentate. the membrane filter at the final concentration for the duration of the run. all costs associated with the size of the plant will go down. In a batch process. and the last stage from 7 to 10%. This means that even with flux decay. the first stage might concentrate from 1 to 2%. Thermodynamically. A new batch must be charged to the reservoir to continue. It is unwise to use the feed and bleed mode to concentrate a batch of solution. and longer membrane life due to greater flux stability. Eventually.g. the second stage from 2 to 4%. The optimum design point will depend on the current cost of power and membranes (along with their productivity). Higher recirculation rates will result in greater productivity (flux). Obviously. Once steady state is achieved. As the permeate is removed. Figure 3. the optimum cartridge flow rate can be different for the later stages operating at high concentration than those operating at low concentration. the first stage is not fully gel polarized (due to the low concentration) and high cross-flow velocities are not as critical as high operating pressures (approaching the threshold pressure). etc. less membrane area. For example. Often.216
Handbook of Industrial Membrane Technology
In a batch operating mode. Since flux decreases with increasing concentration.. unless a purge stream is taken from the loop. the retentate is circulated back to the feed reservoir.) Though the optimum number of modules per stage will vary. the additional pumps and controllers for each stage add to the capital cost and must be considered in arriving at an optimum number of stages. The feed and bleed mode permits continuous filtration. The concentration of retained species will continue to increase with ever decreasing flux.81 shows some of the major operating costs. the concentration ratio and recovery will remain constant. the volume in the reservoir becomes too small to pump and the run is over. the bleed stream of the first stage becomes the feed stream of the second stage. They also require larger pumps and increased power costs.
. Increasing the crossflow rate through a module decreases the membrane replacement costs due to less area (higher flux) and longer life (greater flux stability).

In some applications. UF of electrocoat paint). and the objective is to re-
. Further.g. fractionation of polymers and proteins).. In other applications. gallmin
Figure 3.81: Determination of optimum recirculation rate. the product is the retentate. and the objective is to concentrate or purify the retained species (passing unwanted contaminants through the membrane).Ultrafiltration
217
oo-
Labor t supervision t maintenance
I
I
I
1
1
O'O
20
40
Cartridge
60
60
100
flaw rate. many of the projected applications have not yet materialized (e.g. the product is the permeate.
APPLICATIONS The largest current applications for UF (in terms of installed membrane area) were not envisioned in the early sixties (e..

Further. The economic justification for UF is usually found in improved IC yields due to better quality water. Traditionally. resin fines and other particles. the application of MF and RO in a deionized water loop was discussed. and Japanese semiconductor plants is that most of the large Japanese plants are using UF. Water is fed at 30 psig through a 47 mm diameter 0.56 and 2. One significant difference between U. In one case. For example.. will be retained by UF. Recent reports (DM Data Inc. plants have UF. recovery and use of the product will often pay for’the pollution abatement. Typically. UF was justified on the basis of savings in the MF replacement cost since reducing the load on the MF cartridges naturally extends their life. and amortization are included. Currently.58) to remove microorganisms. power usage. manufacturers get. Originally. Ultrapure Water Semiconductor Industry.000 MWCO UF membrane can remove all particles and macromolecules larger than 0. and the need for even greater water purity has prompted the use of UF. MF operating costs are usually no more than 509 per 1.000 gallons. and a 10.22 p membrane filter
.218
Handbook of Industrial Membrane Technology
move unwanted contaminants which are large enough to be retained by the membrane. in caseswhere the plant effluent is at elevated temperatures.005 /.I. UF is being used after the ion exchange units and before MF.This means that high molecular weight organics and colloidal particles. UF operating costs often run more than this when membrane replacement costs. “Pin-holing” in angstrom thick oxides can occur due to pyrogens. An 80. only a few U. which pass through an MF membrane. Even with replacement every other week. Figure 4. MF cartridges can last from 3 to 6 months (even without UF). Arizona) indicate that Japanese manufacturers of silicon chips get almost twice the yield of usable chips that U. Scottsdale. the economic incentive is found in the avoidance of a sewer tax or shut down by the Environmental Protection Agency. Except for ultrapure water plants where the DI water is dirty and the MF cartridges must be changed every two weeks or less. Increasingly. In Chapter 2. In a few applications.002 cc. the hot permeate may be reused in the plant saving additional energy costs.000 MWCO can remove particles down to 0. The effectiveness of UF in particle/colloid filtration can be measured with an MF plugging test similar to the “silt density index” (see Chapter 4.10). The rapid miniaturization of integrated circuits (I&). maintenance.S..S. the savings in replacement cartridges and labor were expected to pay for the U F plant in less than a year. MF operating costs are considerably less than UF. both retentate and filtrate are important. In these cases. if a valuable product or by-product is a pollutant in a waste stream. The clean permeate may then be reused in the plant. MF has been used after the ion exchange columns (see Figures 2. There are only a few applications where UF can be justified for pollution abatement without the recovery of a valuable by-product or energy credit. dead bacteria fragments/pyrogens (down to 10.S. The organics adhere to the silicon chips resulting in poor adhesion of the polymer masks used to control etching. labor.000 MW) can adhere to wafer surfaces resulting in MOS gate failures.

there is usually no problem in meeting the specification cited in Table 2. It should be realized that the water is relatively clean (the UF unit is operating in the single-pass mode with 95% recovery). Figure 3.2 ppm TOC.) The Center for Disease Control has reportedM that spiral-wound modules produce water with consistently lower bacteria counts than hollow-fiber modules when challenged with high populations of bacteria in the feed water.82:
Effectiveness of UF (ultrapure water) using MF-plugging test. all UF membranes must be sanitized periodically. it is not as “absolute” as an MF membrane (see Chapter 2).4 lists common sanitizing agents.
.05 ppm TOC. a 20.0. Table 3. and that the use of a 47 mm filter accelerates the time to plugging.000 MWCO membrane. (Replacement should be infrequent. With an 80. Sodium hypochlorite is preferred. bacteria counts on the membrane increase exponentially.7 (Chapter 2) of less than 0.000 or 10.Ultrafiltration
219
monitoring the flow rate with time. Though the retention of microorganisms by a UF membrane is well over 99%. In some case. For this reason.22 llIC2ONS
loo I TIIIE (HOURS) 5
10 50
100
Figure 3. For this reason.000 MWCO membrane will be necessary to meet the more stringent Japanese specification of less than 0. With continued filtration.
UF PLUO TEST FILTER . It also attacks stainless steel. it is wise to install an MF cartridge at the point-of-use for final insurance.824’ shows the improvement in the quality of the UF permeate at the Western Electric plant in Allentown. PA.
The removal of organics is measured as a reduction in total organic carbon (TOC). but cannot be used with polyamide membranes.

since GNB grow in solutions containing minimum nutrients. If the LPS is reduced to its smallest subunit. they form micelle-like aggregates. Though MF membranes can remove the microorganisms and particles above 0. thus. Pharmaceutical Industry. and granulomas. However. particularly in persons who are ill or whose immune systems are weaker than normal. colloids and organics down to an average size of 0.000 MWCO UF membrane can remove all pyrogens.. these units are operated dead-ended with a periodic fast-forward flush to remove accumulated particles inside the hollow fibers. Indeed. The word “pyrogen” means “fever-producing” and has been used to cover any substance which causes a body temperature increase on injection. In some cases. 47 Experiments show that the 10.Ultrafiltration
221
A more recent development is the use of UF for point-of-use processing. particles larger than 3 &r can block blood vessels resulting in partial occlusion of retina arteries (leaving blind spots). It is crucial that all microorganisms. the plastic (PVC) piping to the various rinse stations collects particles and leaches organics. Further.
. It involved the measurement of the rise of body temperature in rabbits upon intravenous injection of a test product. Water is.005 ~1.000 MWCO ultrafilter but be retained by a 10. The smallest blood vessels in the body have a diameter slightly less than 3 p. UF has also been used in ultrapure water loops before the ion exchanger units to protect resins from fouling and as pretreatment for RO. a pyrogenic response can result from the presence of dead bacteria or even cellular debris. pyrogens can even be entrained and carried along with water droplets into the condenser during evaporation.000 MWCO membrane. Thus.&they cannot effectively remove pyrogens. it will pass through a 1.000 MWCO membrane will reduce the pyrogen concentration by 2 to 5 orders of magnitude. Because the LPS molecules contain both hydrophobic and hydrophilic regions. It is well known that pyrogens can produce shock and even death. pyrogens. in recent years considerable progress4’ has been made in identifying pyrogens as lipolysaccharides (LPSj primarily from Gram-negative bacteria (GNB).000. and other particles be removed from water for injection (WFI). In fact. pyrogens must be barred entrance to the blood stream. When UF is installed in the central syste’m. These materials occur in the outermost layer of the cell walls.4a Other researchers4’ seeking to remove pyrogens from peritoneal dialysis solutions found that a 50. the first compendia1 pyrogen test was published in the 1942 edition of the United States Pharmacopeia. This explains why normal autoclaving does not destroy pyrogens and filtration through 0.2 p sterilizing membranes does not remove them.47 Historically.2 /. the FDA approved the Limulus amebocyte lysate (LAL) test which can be run for less than 10% of the cost of a rabbit test. even distilled water can evoke a pyrogenic response. the most common ingredient in pharmaceutical products. of course.000 MWCO membrane was adequate and provided a higher flux. clot formation and emboli. A small 4 gpm hollow fiber unit at the rinse station removes all particles. Because of their fever-producing capability. pyrogens have been defined on the basis of their physiological action rather than on their chemical characteristics. More recently (1977). It has been repeatedly demonstrated that a 10.

The electrocoat process for primer coating involves the electrophoretic deposition of charged colloidal resinous particles in aqueous dispersion onto a conductive substrate such as an automobile body. all excess paint is recovered alleviating a serious pollution problem and. The process is universally favored because once the paint particles are deposited. metal furniture. thin films. envisioned UF in a closed loop rinse system (see Figure 3.000 UF units in the automotive industry alone-not to mention appliance manufacturing. Rinse
I
Paint makeup
I
To dr. and other metal electrocoat operations. There are over 1. PPG (a paint manufacturer). Further. therefore. It is. Later. coil coating. coherent and defect free coating-even on sharp edges and in recessed areas inaccessible to other methods. This results in the accumulation of several water-soluble ions in solution with consequent decreased “throwing power”. This dilute paint cannot be returned to the tank due to the excess water. a portion of the ultrafiltrate is purged to the drain to remove ionic contaminants.84) to solve the pollution problem as well as the contaminant build-up. at the same time. justifying the UF unit economically. and pinholing due to rupture. The majority of the permeate is used in the rinse tunnel to wash excess drag-out paint back into the tank. lower rupture voltage.oin Figure 3. At first. The continuous feed of paint to the tank is balanced only by the deposition of charged particles on the substrate. Thus. Virtually every automobile plant in the world uses electrocoat for the undercoat and virtually every installation utilizes UF. In this system.84: UF of electrocoat paint. they insulate the body from further deposition at that point. staining. discarded to the sewer resulting in a horrendous pollution problem and tremendous lossesof paint. UF was used as a kidney to remove contaminating ions. The result is an extremely uniform. The impressed electric field thereby causes the migration of paint particles to uncoated areas. in the early nineteen-sixties. single-rinse system. UF provides a cost effective way of solving several problems associated with electrodeposition systems.222
Handbook of Industrial Membrane Technology
Electrocoat Paint The electrocoat paint market constitutes the largest single application of UF in the world.
i
. excess drag-out paint from the paint tank must be rinsed off the metal part to prevent “orange-peel” and other coating anomalies.

86).e..85:
UF of electrocoat paint.2%) from this rinse is used in the first and second rinse stations. The ultrafiltrate is used only at the last rinse station.Ultrafiltration
223
The size of the UF unit may be reduced considerably by utilizing a countercurrent rinse scheme (Figure 3. Additional savings result from reduced deionized water use. multi-rinse system. most electrocoat tanks were anodic. The paint savings alone typically pay for the UF unit is less than six months. lowered waste treatment costs. i. The dilute paint (0.85). the work piece was positively charged and served as the anode while the paint particles were negatively charged.
In the nineteen-sixties. These systems were ideal since many UF membranes tend to be slightly electronegative with good fouling resistance and excellent flux stability (Figure 3. Optional DI rinse -_
1st rinse
2d rinse
3d rinse
To d&n
Figure 3.86: membrane.05 to 0. and better control of bath composition.
UF of electrocoat paint-stability
of anodic paint flux with XM-50
. eventually returning all excess paint to the paint tank.”
Figure 3.

87: UF of electrocoat membrane (positively charged).50 The advantage of the cathodic process is related to the fact that there is no metal oxidation and hence no metal ions in the paint film..... II II IO I2 I4 10 IO 20222420 I a . Further. Chemical coagulation/flotation or contract hauling are usually more expensive alternatives.. floe and trapped water remain in the oil phase... Metal cleaning tanks and alkaline degreasing baths gener-
... a cathodic paint process was developed where the work piece was negatively charged and the paint particles positively charged... Unstable oil-water emulsions can be broken mechanically or chemically. This was a superior coating process but led to severe fouling problems with some of the standard UF membranes (Figure 3.
--*---2 III.amenable to treatment by UF..87 or membranes with less electronegativity.. and Iubricants for machining.. 2 4 III.. Oil-Water Separations
paint-stability
of cathodic paint with CXM
Billions of gallons of oily wastewater are generated daily. The membrane fouling problem was solved by using positively charged membranes as in Figure 3.. all chemical treatment methods produce a sludge in which the dirt. Stable oil-water emulsions are generated in many diverse industries. So22 243628 TIME (DAYS)
SHUTDOWN
PERIOD
6
8
4042444643
805234~6~606264666~
Figure 3.r. It is the stable oil-water emulsions which are most difficult and at the same time most .224
Handbook of Industrial Membrane Technology
In the nineteen-seventies. giving improved resistance of the finished article to corrosion. Industrial oily wastewaters can be divided into three broad categories according to the distribution of the oil phase: (1) Free-floating oil (2) Unstable oil-water emulsions
(3) Stable oil-water emulsions Free oil is readily removed by mechanical gravity separation devices. Strict environmental legislation now requires industry to clean up these wastes.. /Vera/ working operations use water-soluble coolants. cutting and grinding oils.63)..

Ultrafiltration
225
ate an oily wastewater.
. the hot clear. reduces detergent costs. A well run phosphate line counterflows the rinse stages into the first stage cleaner tank.OOD GPD
~~ Make-Up
AFTER AA. Further. Metal cleaning and wool scouring wastes are illustrative of the diversity of oily wastewater treatable by UF.
BEFORE
Work Flow b Make-Up Make-Up Make-Up Make-Up
mm
Stage 1 Detergent Wash
I A +A
Stage 2 Water Rinse \
A
I
+A A
I
+A A
I
+A AA
Stage 3 Water Rinse !?hO . Metal cleaning operations normally precede painting or plating operationsthe objective being to remove dirt and grease. Textile manufacturing will produce natural oils from wool scouring or fabric finishing oils.AA +-Stage1 Feed 2 MakeUP Stage 2 Ultrafiltrate +.88:
Concentrate
Flow schematic of prepaint phosphating line with and without UF. etc. The phosphating treatment prepares the metal surface for bonding with the paint. detergentladen washer solution is then returned for reuse in the first stage reservoir.000 \ a GPD
Stage 6 Neutralizer
Drain $ 40.000 ’ GPD
Stage 6
1c Ultrafiltration System Figure 3. Rolling and drawing operations use oil lubricants and coolants. Peel oil is another example from the citrus industry. Food processing has waste streams with natural fats and oils from animal and plant processing-particularly vegetable oily wastes.-
Make-Up
Make-Up
I
A+A
MakeUP Stage 3 X x Drains Closed Off
A +A
I
A +A
I
AA
Stage 4
Stage 5 \ Drain 30. the savings in disposal charges alone (30-fold reduction in volume) can pay for the UF system in less than two years. UF can be very useful in this application in extending the useful life of the wash water and reducing the waste disposal problem. fresh water is introduced into the third stage rinse which overflows into the second stage. The buildup of oil and dirt in the alkaline cleaning stage is removed by UF. and decreases spray nozzle clean-out maintenance time. Figure 3.88 is a schematic of a prepaint phosphating line before and after introduction of UF. \ G:
Stage 4 Phosphating
Stage 5 Water Rinse Drain 30. This closedcycle counterflow system lowers the oil level in the third stage.

If the oil concentrate can be burned in a boiler.1 to 10% oil in a stable emulsion. In the presence of poorly stabilized oil emulsions. and other disposal methods. UF has been effective in reducing the COD to less than 15. The waste stream is highly polluting with a COD of 80. The secret of successful UF is to maintain discrete and stable emulsoid particles of oil (generally over 0. Low lubricity emulsions (with a high “synthetic” oil content) are usually characterized by high flux and low fouling rates. oil concentrations of over 50% can support combustion. cellulosic membranes significantly outperform medium surface free energy membranes which.I or below). aqueous detergent solutions are used to remove contaminants from raw wool-principally wool grease and suint (sheep sweat) with smaller amounts of soil and fecal matter. oil in the permeate will generally be less than 10 to 50 ppm.01 /. Indeed.1 Oil-water emulsions behave like colloidal suspensions showing typical “gelconcentrations” of 75 to 80% (close packing of spheres) and a flux independent of pressure (in the “gel-polarized regime”). During scouring. The difficulty with free oil or unstable emulsions is that the oil accumulates at the membrane interface and may form a continuous layer which preferentially wets the membrane over water (the surface tension is lower). Oils containing large amounts of “tramp oil” with moderate to low lubricity show lower flux and require more frequent membrane cleaning. the suint dissolves. A limited amount of free oil can be processed but usually quantities above 1 to 5% are removed with a centrifuge prior to UF. burning. (See Chapter 2 on the bubble point test). If the chemical nature of the oily wastewater requires a more chemically resistant hydrophobic membrane. In this case. Cleaning is normally done with detergent solutions at elevated temperatures (60°C) and high pH for 1 to 2 hours.rin size) which are larger than the membrane pore size (0. it is important that a high crossflow velocity be maintained.
. UF can concentrate the oil up to 40 to 60%.S. Despite the fact that these are among the highest quality products refined from our best crude oils.226
Handbook of industrial Membrane Technology
In the wool scouring process. This can reduce the contract hauling costs by two orders of magnitude.1 /. Oily waste waters suitable for treatment by UF contain 0. significantly outperform low surface free energy membranes.25 billion gallons of oil every year by changing the oil in their automobile crankcase. and the mineral particles become suspended. Reclamation of Waste Lubricating Oil U. motorists discard about 1. Usually. (Oil wetting of hydrophobic membranes begins at stagnation points. the membrane will pass oil and retain water.. Noncellulosic membranes can become wetted with oil more easily-losing their water flux and rejection for oil.000 ppm. only about 100 million gallons reach recyclers to be re-refined into clean lubricants. When this is the case. High lubricity oils (containing natural fatty materials) exhibit low flux and are prone to foul the membrane. any fuel credit can help offset the cost of the UF plant. The rest is burned as fuel or disposed of as waste. the wool grease is emulsified by the detergent solution. a conservative estimate is that over 500 million gallons of used lubricating oil are annually injected directly into the environment through landfills.000 ppm as well as decreasing disposal costs. Further. in turn.

Ultrafiltration

227

The idea of recycling these lubricants has flourished when oil is scarce. Presently, the oil glut has put many re-refiners out of business. In 1935, Adolph Hitler prodded German industry to collect and recycle its dirty lubricating oil to help reduce oil imports. Since then, subsidies and legislation have made the concept a permanent fixture in Germany. It is said that even today some twothirds of West Germany’s waste crankcase oil is recycled. Eventual shortages of petroleum reserveswill necessitate some form of recycling. The standard re-refining process uses sulfuric acid to dissolve sludge. The oil phase is further cleaned by clay filtration. The process is so inefficient that one-third of theoil is lost and disposal of the acid solution is increasingly difficult. The “water-alcohol“ method utilizes isopropyl alcohol to decompose metallic soaps. The mixture is then centrifuged to yield clean oil and a watery metallic sludge. Alcohol is recovered for reuse. UF has the potential of removing all contaminants, but the flux is very low at room temperature. Inorganic UF membranes operating at 300°C and a pressure of 7 bar (105 psi) are capable of processing the oil economically. In one plant in Europe, where the spent lubricating oil is pretreated with thermal shock and centrifugation at 18O”C, the UF flux is reported stable between 1000-2000 LSMD (25-50 GSFL) over six months without cleaning. Decontamination of Crude Oil Nearly all metal contained in crude oil is either chemically bound with asphaltene type structures or associated with asphaltenes (IOOO-10,000 daltons). It is necessary to remove these metals to prevent contamination of catalysts. The present process uses a solvent like hexane or pentane to flocculate the asphaltenes and is relatively expensive. UF can retain the metal asphaltene complexes while passing virtually all other crude oil fractions. Again, the flux is low due to high viscosity. However, an inorganic membrane operating at 330°C on a 10% asphaltene feed produced a permeate with less than 0.5% asphaltenes. The removal of the carbon-forming asphaltenes and catalyst-fouling metals also reduces the process costs for desulfurization of fuel oil. PUA Recovery Polyvinyl alcohol (PUA) is used as a sizing agent to improve the strength and surface characteristics of warp yarns prior to weaving operations, where they are subjected to considerable abrasion and tensile stress. The traditional warp sizing agents (starch, gelatins and gums) have been replaced by improved synthetic materials like PUA or sodium carboxymethylcellulose (CMC). PUA sales for this application now exceed 40 million pounds annually. Later in the textile manufacturing process, the sizing agent is removed from the cloth by scouring before dyeing and finishing operations. Because the synthetic sizing agents are not biodegradable, they pose a serious pollution problem. Although chemical precipitation of these sizing agents is possible, there is a sludge disposal problem. Due to the high cost of PUA and CMC, their recovery with UF and reuse is a much more attractive solution. Figure 3.89 is a schematic of the PUA recovery system. Not only can the PUA be reused, but hot permeate (185’F) can be recycled directly to the desizing bath. The dollar value of the PUA and energy recovered will generally pay for the UF system in 10 to 18 months.

further

I
3ving ther ratio ins -

processing ,20-50

pm cartridge

filter

i -Makeup water

Ultrafiltration

sys rm LP

-_ -.. __ _..-.. -.. - _. _.- .. Balance tank

Permeate return

I

I

em-.. -..-.. -.. -..

-.. _. _

_ .

-..-..

-..

Slasher
Figure 3.89:

Mixing tank Flow schematic of PVA recovery system.

Ultrafiltration

229

The UF membranes are protected from yarn fibers by a vibrating screen filter and 20 to 50 /J cartridge filters. This makes possible the use of spiral wound modules which have a life of 24 to 30 months. The desizer waste effluent usually contains between 0.5 and 1.5% PVA. The UF concentrates this PVA up to 10% for direct use in the slasher. This final concentration is monitored and automatically controlled by an in-line refractometer. A small amount of desizer waste is purged to drain to prevent the buildup of low MW solutes. Dyestuff Recovery and Purification Many dyes are too small to be retained by a UF membrane. Exceptions include polymeric dyes and indigo; the latter has a low molecular weight (262 daltons) but can be retained by a 50,000 MWCO membrane, when in the oxidized state due to its insolubility in water. The economics of an indigo dye house can be improved dramatically with the recovery of indigo from the plant effluent. Losses through the rinse system can account for more than 10% of the mills’ total indigo consumption. A membrane life of 3 years is reported. Whether recovered by UF or RO, it has been established that recovered dyes can be reused in dyeing operations with no attendant problems and meeting all color specifications.‘l UF has also been evaluateds2 as a method for purification of polymeric dyes made by the attachment of an azo-chromophore onto a polymeric backbone. This is particularly important when these dyes are used as food colorants. The low molecular weight species and oligomers must be removed to ensure that the product will be nonabsorbable following ingestion. As expected, these polymeric dyes were rejected completely by the membrane. Surprisingly, Sunset Yellow (MW 452), Schaeffer’s salt (MW 246) and tartrazine (MW 534) chromophores were significantly rejected by membranes with MWCO’s in the range of 10,000 to 50,000. It is known that dye molecules of this type can associate in solution; but an increase in NaCl concentration, expected to increase the degree of association, decreased the rejection of the dyes. Therefore, it is more probable that the dye molecules adsorb to the membrane making an effective anionic membrane rejecting the charged dyes. At high ionic strength, the charge on the membrane is screened allowing passage. It was found by Cooper et aIs that a sodium chloride concentration of 15 g/2 was necessary to achieve good transport of the small MW chromophores. Alternatively, a high pH’? had a similar effect. Latex Concentration/Recovery The concentration of latex emulsions was one of the early applications suggested for UF. Unfortunately, not all latexes are amenable to processing with UF. Many latexes are unstable under the high shear induced by pumps or even in the thin channels of UF modules. However, the use of diaphragm pumps and careful control of hydrodynamic shear within the module helps prevent coagulation. UF has been used in place of evaporation to concentrate in-process latex streams from 30% to as high as 65% with reduced energy consumption. Polyvinyl chloride is the most prominent latex in this application.

230

Handbook of Industrial Membrane Technology

UF has also been used to clean up polymerization kettle wash waters before disposal. The dilute latex can be concentrated from 0.5 to 25%, thus reducing the volume to be hauled away to ~/SO the original volume. In some cases, the of waste latex is recycled for reformulation. Where there is a significant sewer tax, UF is an economical alternative even without recovery of the latex. The dilution of latex during polymerization kettle rinse-down may deplete the surfactant on the polymer particles or introduce multivalent ions, resulting in decreased latex stability. Adjustment of pH or addition of surfactant can prevent coagulation of the latex on the membrane. If the feed is preconditioned properly, the UF flux is often quite stable. One thousand hours of continuous operation between cleanings is common, When flux decay does occur, detergent washing is usually sufficient to restore flux. In some cases, polymer solvents may be required. Proper selection of a solvent resistant membrane and/or solvents which will dissolve the latex but not affect the membrane is crucial. For PVC latex, the solvents of choice are methyl isobutyl ketone (MIBK) and methyl ethyl ketone (MEK). Styrene butadiene rubber will swell in MIBK, MEK or toluene. Polyvinyl acetate will dissolve in the low MW alcohols such as propyl alcohol. Generally, the membranes are first washed with water, then detergent, followed by another water flush. The system is then drained of all water since it will affect polymer solubility in MEK. Finally, a solvent rinse is employed. If the module is tubular, sponge balls will enhance cleaning. Removal of Heavy Metals UF can remove metal values from metal plating wastes provided the metals are first precipitated to form a colloidal suspension. The process is described in Chapter 2 in the section on cross-flow applications. Even though MF membranes provide a higher flux, they are more prone to flux decay than UF due to internal pore fouling. The long term UF flux may be higher than that from MF. Pulp and Paper Waste Treatment Although UF cannot recover the low MW sugars in waste from the pulp and paper industry, it can remove lignin compounds and most of the color bodies. Pulp is made by separating wood into individual fibers. This is done mechanically or chemically. Only 50% of the material in the tree is used as pulp. The remainder is often discarded as waste in the spent pulping liquor. In the mechanical process, logs are pressed against a rotating grindstone; the resulting fibers are mixed with water and screen filtered. This mechanical pulp contains lignin, gums and mineral salts in addition to the wood fiber. In the chemical pulp process, wood chips are digested with a chemical “cooking liquor” under high pressure and temperature. There are three major cooking processes. The sulfate or kraft process is the. most prominent; about 75% of the wood pulp produced in the U.S. is made by the kraft process. It uses a solution of NaOH and Na2S, called “white liquor,” to remove the lignin that binds the cellulose fibers together in the wood. The loose fibers are then separated from the spent cooking solution-called “black liquor”. The soda process uses caustic (NaOH) and the sulfite process uses sulfurous acid (HsSO,). If the pulp is to be used in high-grade white paper, it is bleached with chlorine.

Ultrafiltration

231

Much of the BOD contribution is pulping wastes comes from low MW carbohydrates. RO can often reduce the BOD by 70 to 90% whereas UF effects only a 45 to 55% reduction in BOD. 53 However, UF can remove 85 to 98% of the color bodies. UF is finding some utility in recovering lignin compounds-used in the production of vanillin, adhesives, detergents, and dispersants. Separating the sugars and lignin compounds solves the problem of sugar contamination in the feed stream to a vanillin plant. Unfortunately, a clean fractionation between the two is not possible even with UF since some of the sugars are bound in lignin-carbohydrate complexes. Table 3.5% shows the retention of sodium base spent sulfite liquor by a series of UF membranes. The multistage UF was carried out in succession from the largest to the smallest pore size-i.e., the permeate from the first stage became the feed for the second stage. Although 90% of the sugar passed through membranes between 100,000 and 20,000 MWCO, the retention increases for 10,000 and 500 MWCO membranes up to 36%. It has been estimated that about ‘16of the sugars are bound in complexes with molecular weights above 10,000 and an additional 10% in complexes above 500 daltons. In other words, ‘/4 of all the monosaccharides present in spent sulfite liquor are bound. On the other hand, 100,000 MWCO membranes retained over 50% of the lignosulfonates. It appears (from Table 3.5)% that about 70% of the lignosulfonate compounds have a MW over 50,000; only 7.2% had a MW less than 10,000.
Table 3.5: UF of Sodium Base Spent Sulfite Liquor-Retention Lignins Original sugar, % Ultrafilter Nomlnal retentlon, mol.wt. Passing Retained Original lignin. % Passing Retained of Sugars 81

It has also been shown that hydrolysis of the spent sulfite liquor with 6% H3P04 can improve lignin retention by raising its MW through polymerisation.s4 The discharge of colored effluent is regulated primarily for esthetic reasons, but it may also interfere with plant and animal life cycles by blocking sunlight. Color bodies from pulp and paper mills are difficult to remove by conventional sedimentation and biological treatment; it requires massive lime treatment and is ineffective with spent sulfite liquor (less than 80% color removal). UF can usually remove over 90% of the color bodies and often achieves over 98% removal. The rejection data of Table 3.6” indicate that more than 97% of the colorbearing material has an effective MW greater than 10,000. In these data, the premeate from the UM 10 membrane was fed to the UM 2 membrane and its per= meate was fed to the UM05 membrane. Thus, the UM2 membrane can remove 99% of the color, 64% of the COD and 43% of the TDS.

Though membranes have been operative in the pulp and paper industry since 1970, growth has been slow. Economic justification has sometimes been marginal; fouling problems often limit membrane life to less than a year. Dairy Applications Cheese Whey Protein Recovery. Perhaps the best publicized application for UF is in cheese whey processing. “Cheese whey” is the supernatant liquid produced in the cheese making process after precipitation of casein from milk. There are two types of whey; “sweet” whey (minimum pH of 5.6) results when rennettype enzymes are used to coagulate the casein to form Gouda and Cheddar cheeses; “acid” whey (maximum pH of 5.1) results from acid-induced coagulation of cottage cheese or casein. The typical composition of cheddar cheese whey and cottage cheese whey is shown in Table 3.7. Whey contains half of the milk solids; i.e., two gallons of milk (17 pounds) yield one pound of cheese and one pound of whey solids. Most of the lactose winds up in the whey along with 20% of the protein. Table 3.7: Composition of Cheddar and Cottage Cheese Whey

Components Total solids Lactose and lactic acid

From Cheddar Cheese % 7.31 5.2. 0.87 0.53 0.71

From Cottage Cheese % 6.53 4.39 0.86 0.61 0.67

protein (amino acid ni¶xogen) Ash Fat

The lactose is the prime contributor to the high BOD of the whey stream (35,000 to 55,000 ppm). The 150 billion pounds of cheese whey produced each year has become a significant pollution problem choking out aquatic life in

Ultrafiltration

233

many streams and lakes and imposing a severe added burden on sewage treatment facilities (1,000 gallons = the waste from 1,800 people). The valuable component of cheese whey is not the lactose but the whey proteins, primarily lactalbumin. The amino acid profile of these proteins is superior nutritionally to casein and is equal to or better than whole egg protein. The heatdenatured form of these proteins has been manufactured for many years usually by heating the cheese whey to precipitate the proteins. The product was tan colored and completely insoluble. With the advent of UF, these proteins could be recovered, concentrated and demineralized athermally. The result was a “whey protein concentrate” (WPC) with improved solubility and other functional properties (emulsification, foamability, water binding, gelation and cloud stability). Unfortunately, UF of cheese whey does little to help the pollution problem since the lactose passes through the membrane. Initially, a two-step process involving RO, after UF, to recover lactose was envisioned (see Figure 3.90). The recovery of proteins in the first stage (UF) and lactose in the second stage (RO) is shown in Table 3.8.s6 The first stage can be justified economically with the value of the high-grade proteins.

Unfortunately, lactose cannot demand a price sufficient to pay for the second stage. Current world production of lactose is less than 5% of that in whey and future demand shows no significant increase. The growth of UF for this application has been hindered by the difficulty of using the permeate profitably. This boils down to a profitable utilization of lactose. Figure 3.91 shows the principal uses currently envisioned” for UF permeate.
WHEY PERMEATE

Pag feed

E
c +Alcohol bother (eg extroctlon

t
or wth other moterialsl hck to cattle urea ammomum lactate I Methone for fuel INDUSTRIAL

Oryinglstngly Crystolltzotaon wth

I
ANIMAL USE

Reactton

Reactcon Biomass

to gwe

+Fermentataon

+Anoeroblc

t orgomc chemicals loctlc ocld. ontiblotlcsl

*Lactose

DIRECT HUMAN USE s~gool~e/Glu~~~~g

I

*Hydrolysis I+ deionizotton + concentrotlon

1
Confectionery Syrup I

Figure 3.91: Principal uses envisioned for UF permeate.

In Switzerland, 92% of the whey is used for pig feed. Reaction with urea to form lactosyl Urea permits greater quantities of non-protein nitrogen to be used in ruminant feeds. The lactose can also be fermented under aerobic conditions to yield a biomass for animal feed with a crude protein content of about 45%. Several commercial plants in France and Austria are producing biomass in this way, but with marginal economics. Anerobic fermentation can be used to produce methane to fuel the steam boilers in the creamery, or to produce alcohol for fuel. One ton of lactose will produce a half a ton of alcohol. In 1983, the first fully integrated commercial plant for converting whey to fuel-grade ethanol was constructed in Manteca, California. It has the capacity to process nearly 500,000 lb of whey per day producing 400,000 gal of anhydrous alcohol. In addition to ethanol, the plant was designed to produce whey protein animal feed in wet cake form and a high-protein liquid feed supplement for cattle, hogs and poultry. Whey wine is an experimental product being test marketed. After deproteinizing with UF, the lactose is fermented for about a week by special lactose-fermenting yeast. The finished product is a pale-yellow, tart, dry wine with a subdued aroma and bouquet.

Ultrafiltration

235

Hydrolysis of lactose to galactose/glucose enhances its solubility (from 22 to 60%) and sweetness (from 15 to 70%). In 1983, a plant utilizing immobilized lactase to hydrolyze deproteinized lactose (UF) to galactose and glucose was constructed in Winchester, KY (Nutrisearch-a joint venture between Kroger and Corning Glass). A similar plant was constructed in Cheshire, England (joint venture between Corning Glass and the Milk Marketing Board). The resulting syrup is used as a sweetener and as a growth medium for baker’s yeast. Figure 3.92 is a process schematic of the world’s largest cheese plant (at time of startup in 1985) in Corona, California. The plant utilizes many of the processes described above producing 5 million lb/year of WPC (by UF) (50 to 75% protein) and 2.2 million gal/year of ethanol for gasohol. The lactose permeate from the UF unit is concentrated by RO before use as a fermentation medium to produce alcohol.

Perhaps the greatest impediment to the use of UF is membrane fouling with a decline in flux; it is particularly severe with cheese whey. Usually a cheese plant cleans its equipment once a day for sanitary purposes with high and low pH. Early cellulose acetate membranes could not withstand the standard cleaning cycles. Today, however, polysulfone and polyvinylidene fluoride membranes may be cleaned with hot (6O’C) solutions pH 1-12. In some cases, nonionic detergents (0.1%) are used to remove fat deposits and enzyme detergents are used to hydrolyze protein deposits. Complexing agents such as EDTA and SHMP are useful in removing inorganic deposits such as insoluble calcium slats. Sanitation for all but the polyamide membranes is accomplished with solutions of sodium hypochlorite (with 100 to 200 ppm available chlorine). Various pretreatments are sometimes utilized to prevent fouling. We have already seen that pH has a marked affect on flux (see Figure 3.37) by altering the solubility of the proteins; pH adjustment can also be beneficial in preventing precipitation on the membrane. Some of the best work on fouling and pretreatments has been done by Lee and Merson at the University of California at Davis.58r5g.They examined membrane deposits with a scanning electron microscope in an attempt to develop preventive measures against membrane fouling. Systemic prefiltration experiments were conducted6’ to study the effects of removing certain constituents from the whey. The hypothesis was that small amounts of relatively largesized protein structures and microorganisms provide anchor points which facilitate attachment of the lactoglobulin proteins. Prefiltration designed to remove these anchor structures did indeed lead to a higher flux with greater stability. For example, passing untreated whey through Whatman #I filter paper doubled the permeation rate of a 10,000 MWCO membrane. They also found that calcium sequestration and raising the ionic strength of the whey increased the flux. It is expected that new surface treatments of the membrane itself will provide significant improvements in flux stability. UF should be a standard whey processing step in the future. Even today it is estimated that UF treats 10 to 15% of the total world production. Milk Concentration. One of the most promising innovations in cheese making is to concentrate the casein and proteins in milk with UF prior to coagulation of the casein. Maubois61 is the inventor of this process. Though the UF permeate still contains lactose, all of the proteins are incorporated in the cheese resulting in a 5 to 20% reduction in milk usage depending on the type of cheese.62 A two-fold concentration can double the production capacity of each cheese vat and reduce rennet consumption by up to 80%. The savings in milk usage alone can often pay for the UF plant in less than one year. At first, UF was used to make soft cheese of the Camembert type. Next, semihard cheese like mozzarella was made and more recently, hard cheeses like Cheddar. In 1981, 105,000 tons of hard cheese was made by this process. Denmark has taken the lead in the production of cheese (primarily fetal by UF; in 1981, 92,000 tons (32% of Denmark’s total cheese production) was produced by UF. Because some of the minerals are lost with UF, there are some differences between traditional cheese and that made with UF. In the U.S., this application has been hindered by concern over whether cheese made with UF meets the established standards of identity. In the Spring of 1985, the FDA said feta cheese made from UF may be acceptable in the U.S.A.

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237

There is economic incentive to concentrate the milk on the dairy farm rather than at the cheese factory. This reduces hauling costs but requires a greater capital investment in automating a number of small UF units on many dairy farms. Pioneered in France by Maubois, the first U.S. commercial study on-the-farm was initiated at a 900 cow farm near Lodi, CA in late 1984. The UF project is supervised by R. Zall of Cornell University and is co-sponsored by the California Milk Advisory Board and Dairy Research Inc. Raw milk is heated to 163’F to stabilize micelles in the casein fraction and seal the whey proteins. The milk is then concentrated by UF to half of the original volume and is shipped to several cheese makers producing Monterey Jack, Cheddar and mozzarella. The permeate (containing lactose and soluble salts) is fed to the cows. Preliminary reports indicate significant savings due to reduced transportation and refrigeration costs plus a dollar credit for livestock feed. If fouling of UF membranes is a problem with cheese whey, it is an even greater problem when concentrating milk. The units on the Lodi farm above are cleaned as follows: (1) Flush with permeate for 5 minutes

A considerable improvement in UF flux (+lOO%) and stability can be achieved by prefiltering the milk with a 0.4 /.4pore size MF membrane to remove bacteria, fat, lipoproteins and coagulated (denatured) proteins. The MF must be operated in cross-flow to prevent plugging. There are other markets for concentrated milk besides cheese making. Condensed milk or dietary milks (without lactose or salts) can be easily prepared with UF. The solids concentration can be increased to well over 20% but with large increases in viscosity (over 30,000 cp). Food and Beverage Applications (Non-Dairy) Soy Whey. The pollution problem facing cheese makers is also a problem for food processors isolating protein from soybeans, cottonseed and other oilseeds. Purification of soy protein from defatted soy flake involves extraction and precipitation. The supernatant after precipitation is similar to cheese whey except that the protein concentration is only one-third to one-half of that in cheese whey. These low protein concentrations reduce the incentive to use UF; as a result, very few plants use it. Egg White. The wide spread use of egg white in the baking and candy industries (more than 300 million pounds in the U.S. per annum) arises from its ability to form stable foams which can support relatively large quantities of sugar and/or flour. Fresh egg white contains approximately 88.5% water; it is desirable to remove some of this water to save processing, packaging, refrigeration and transportation costs. However, the proteins are sensitive to heat. Concentration by evaporation or spray drying destroys the functional properties of the

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proteins by thermal denaturation. Athermal concentration by UF is ideal. UF is more appropriate than RO for this application because glucose and inorganic salts pass through the membrane. Glucose reacts with the amino acids in the proteins (the “Maillard reaction”) resulting in the development of caramellike flavors and odors, evolution of CO* and the eventual separation of an airinsoluble, dark brown, humus-like material.56 Several alternative methods have been developed to retard this browning reaction. The most prevalent method is to reduce free glucose by a controlled bacteria or yeast fermentation to gluconic acid. The fermentation takes anywhere from 2 to 72 hours and adds considerable expense. With UF, protein concentration and glucose reduction can be accomplished simultaneously. UF concentrates have been shown to have equivalent functional properties with fresh egg white. ” Even dried solids obtained by spray drying the UF concentrate show improved characteristics over that using spray drying alone.63 Economic analysis64 indicates that savings in transportation and drying costs result in a 40% return on the capital investment required for the UF plant. Gelatin. Gelatin is a colloidal protein substance whose principal value depends on its coagulative, protective and adhesive powers. It swells up and absorbs 5 to 10 times its weight of water. Water containing only 1% gelatin by weight will gel when cold. There are threedifferent typesof gelatin: (I) edible, (2) photographic and (3) inedible (i.e., glue). All types are derived from hydrolysis of collagen (MW 100,000). Since the hydrolysate contains only 3 to 15% gelatin, it must be concentrated. The traditional method has been steam evaporation followed by drum dryers. Again, these elevated temperatures degrade the product with a resulting loss in gelling power. In addition, the salts in crude gelatin tend to degrade the polymer structure and are commonly removed by ion exchange. UF can perform the required concentration at lower temperatures with simultaneous removal of salts.65 It is of course true that even with UF elevated temperatures (above 5O’C) must be used to keep gelatin liquid. It is advantageous to operate at 70°C to take advantage of the lower viscosity and higher flux. Fruit Juice. Apple juice processors are using UF to increase juice recovery, improve clarity, reduce materials cost, and to cold sterilize for aseptic packaging. Before the advent of UF, juice precessors hydrolyzed cloud-stabilizing polysaccharides such as pectin and starch with pectinase, coagulated with gelatin, and then clarified the cloudy cider with a diatomaceous earth filter. The materials collected on the filter provided an ideal substrate on which to grow and multiply bacteria. This resulted in high microbial counts in the juice, but high temperature pasteurization is not desirable because of its effect on flavor. UF can remove the colloidal haze in apple juice (due to pectins, starch or protein) while all the sugars (fructose, glucose, sucrose) and flavors (sorbital, esters, organic acids) freely permeate the membrane. In the process, all bacteria are removed to produce a clear “sparkling” apple juice which has a long shelf life. Table 3.9 is a comparison of operating costs for a 70,000 GPD plant replacing the conventional filter press with UF. The major savings are associated with the elimination of diatomaceous earth filtration and the reduction in enzyme (pectinase) requirements. Even with amortization, the operation costs with UF are only 45% of the conventional system costs. The ROI is 42% with a payback in 2.4 years.

carbohydrates and proteins normally removed by diatomaceous earth and EK pads after fining. while conventional juice after DE filtration ranges from 1. Taste panels have been surprised to find that a UFprocessed wine had cleaner nose and better body.
. Often. this reduces fouling in the RO module and permits a higher RO flux.000 MWCO membrane was chosen to filter a Sauvignon Blanc after the taste panel noticed a significant difference in body with a 1. winemakers have been reluctant to try even a 0. precipitation is not necessary.6s a 10. The UF permeate will typically read from 0. the UF concentrate may be diafiltered (see the section on pharmaceutical applications) to recover the sugar left in the concentrate.240
Handbook of Industrial Membrane Technology
In some cases. In spite of DE and pad filtration. Membrane life for tubular systems appears to be about 9 months. colloid-free filtrate from which sucrose can be directly crystallized in high yield and purity. UF produces a sterile permeate of excellent clarity. tannins. UF can remove the yeast and/or bacteria normally removed by MF. Bentonite and gelatin are added as fining agents to precipitate colloidal proteins so they can be removed with conventional filters.5 to 3. MF can clog within 4 to 5 hours without UF.45 /. Cranberry juice and other juices can be clarified in the same way.r and above) have been used for years to cold sterilize wine and to remove particles and haze. gums and pectins. Some wineries have tried UF as a last resort when wines show recurring heat stability problems. UF of raw sugar juice yields a clear. The current interest in UF is primarily because of its ability to replace several processing steps with a single filtration. Beer.6 NTUs on a turbidimeter. The haze in wine is often due to heat-unstable proteins.0 NTUs. Wine. In addition. UF can remove these color bodies along with the slight bitterness from phenols. There are two primary applications for UF emerging within breweries: treatment of tank sediment and alcohol reduction. UF produces a superior product with a shelf life of up to 24 months with “sparkling” clarity.4 to 0.000 MWCO membrane did not provide the heat stability desired.67 In some cases.000 MWCO membrane and the lab determined that a 100. With UF. Press wines are sometimes slightly brown due to higher levels of oxidized phenolics. MF is kept as a final filter for insurance. in the past. Product yields with UF are typically 96 to 99% compared to 80 to 94% without. UF prolongs the life of this filter indefinitely. A number of different kinds of filtration have been used to extract beer from tank sediment-including yeast presses and DE filtration. the UF permeate is then concentrated by RO. MF membranes (0. UF can also be used to increase the tannin content in red wines with poor body or mouth feel (tannin is retained by the membrane). As a rule. the selection of the best MWCO has been determined by the taste panel. several wines processed with UF have recently won awards as some of the best wines in their class. oxidized/polymerized phenols. Yet. The recent interest in UF of wine is surprising. UF appears to accentuate acidity and the lighter fruity elements while eliminating tannic-related “off notes”. The beer recovered with conventional filter techniques is not of good quality. In one case. It can remove all colloids.2 /J pore size for fear it would affect the taste.

Broadly defined. For example. In “diafiltration”. Today’s new biotechnology focuses on applications for agriculture and industry. vitamins. To date. Beer fermented to a low alcohol strength does not develop a satisfactory body. taste and flavor (too little wort). brewing. In this sense. the product is in the retentate. One of the fastest growing markets for UF and chromatography is in biotechnology. the permeate volume is replaced by an equal volume of pure water so as to not change the concentration of the retained components in the beer. “old biotechnology” would produce wine by fermentation but “new technology” seeks to use r DNA to produce wine with a higher alcohol content. polysaccharides. hormones. cellular components. Most industrial biological syntheses are carried out on a batch scale. vaccines and antibiotics are all processed with UF. with a small amount of product recovered from large quantities of water. In some cases. which is why we are considering pharmaceutical and biotechnology applications together. fractionation and sterilization are all practiced in one form or another. This reprograms the microbe to perform differently and creates a hybrid without the time delays and chance which occurs in cross-breeding of plants and animals by traditional means. the value of the materials recovered is not high. Occasionally. and cheese making. Genetic engineering can now take a gene from one organism and slip it into the genetic material of another organism. Proteins. most applications in industry are part of the pharmaceutical industry. as the water-alcohol mixture passes through the membrane. in others. The commercial success of r DNA and cell fusion technologies is dependent on advances in “bio-process engineering”. This makes possible the synthesis of new hormones. The plant must be justified on the basis of pollution abatement. biotechnology has been practiced by man since the dawn of civilization in baking. and wastes. “New biotechnology” focuses on the industrial use of recombinant DNA (r DNA) and cell fusion. Corn Starch. both retentate and permeate contain valuable products. viruses. UF can recover colloidal starch and other high molecularweight compounds that contribute to the high COD level of the waste effluent from a corn starch plant.67 Pharmaceutical and Biotechnology Applications The diversity of applications for UF in the pharmaceutical industry is unequaled in any other industry group. to improve plants or animals. biotechnology includes any technique that uses living organisms or parts of organisms to make or modify products. The fact that all of these operations are possible with UF at ambient temperatures without phase change or addition of chemicals/solvents makes it an ideal separation tool for labile drugs and biologicals. Concentration. it is in the permeate. antibiotics. the alcohol can be reduced to any desired level from the same type of beer.Ultrafiltration
241
There is an increasing demand for low-alcohol beers which have the beer taste but with fewer calories and not enough alcohol to cause drunkenness. Using UF in the diafiltration mode (see the section on pharmaceutical applications). nutrients. desalting. Unfortunately. insulin and interferon. Even the water to make up these pharmaceutical solutions is often sterilized and depyrogenated by UF (see the discussion on ultrapure water in an earlier section of this chapter). purification. It is in this
. or to develop microorganisms for specific uses.

UF systems can be designed to recover enzymes athermally with minimum shear inactivation (using diaphragm pumps). From a chemical engineering point of view they are the ideal catalyst. The molecular weight of enzymes varies from 12. As discussed in the section on fractionation of solutes. oligosaccharides and salts. the application of cross-flow MF to cell her_ vesting was discussed. The number of existing enzymes is estimated to be more than 10. On the other hand. UF is applied as above. germination seeds. and affinity chromatography (reversibly binding some proteins with very specific ligands/antibodies) can fractionate proteins with a precision unknown in current UF technology (see the section on fractionation of solutes). Ion exchange chromatography (separating on the basis of electrical charge).242
Handbook of Industrial Membrane Technology
recovery and/or concentration and purification of products where membranes may be most useful.OOO. high performance liquid chromatography (Using gel exclusion principles depending on the pore sizes of the beads). In Chapter 2. The reason is that asymmetric UF membranes are less prone to fouling and internal pore plugging than are MF membranes. fermented bran. The two most important separation techniques for biotechnology appear to be membranes and chromatography. After extraction and prefiltration to remove suspended solids and particulates. it is preferred over MF for cell harvesting. filter presses. After separation from the biomass with open membranes. PH end ion composition (“Cohn precipitation”). Cell Harvesting. However. Membrane separations are more easily scaled UP and more amenable to continuous processing at less expense than chromatography.000 of which more than 100 have been purified in crystalline form and over 600 in fairly purified form. chromatography can separate proteins and other macromolecules with greater specificity.700 (ribonuclease) to over l. All enzymes are proteins. conjugated proteins or metalloproteins containing one or more active sites per molecule. Where the products of the fermentation can pass through the UF membrane. Inorganic catalysts are seldom as specific (enzyme catalysts yield almost no by-products) and require high temperatures and/or extremes in PH to be effective. Purified therapeutic proteins are derived from human blood plasma via cryoprecipitation followed by sequential precipitations effected through increasing ethanol concentrations at controlled temperature. enzymes may be inactivated temporarily or permanently by heat. or centrifuges. UF cannot be used to fractionate the various plasma
. Blood Plasma Processing. The majority of microbial enzymes are produced in relatively small batch fermenters. UF is used to concentrate the enzyme and remove small peptides. or shear forces. d-carboxylase). Because of their protein structure. or the “beer” from the submerged fermentation of selected strains of microorganisms. Enzymes are highly specific biological catalysts which have an optimum pH and temperature (usually fairly mild conditions) for maximum activity. Enzyme Concentration and Purification. antibodies and antigens have been immobilized on both MF and UF membranes and are sold as immuno affinity membranes for immunodiagnostic tests. Crude extracts of enzymes are also obtained from macerated tissues. These membranes have the potential of combining the specificity of affinity chromatography with the low-cost processing of UF.OOO (L-glutamate dehydrogenase. chemicals.

Indeed. (Notice that even though salt is removed in the permeate. In this way. (As a consequence. Albumin preparations are routinely concentrated with a 10.94 is an example of the discontinuous process occurring in three steps: (1) concentration. Further. there is less denaturation. it is particularly susceptible to thermal and shear denaturation. UF has also been used to concentrate antihemophilic coagulation factor VIII (AHF) with no measurable loss in activity. the concentration and purification of the various protein fractions was among the earliest applications for UF in the pharmaceutical industry. no latent heat is required. the enzyme is reconcentrated three-fold from 6 to 18%. from 18 to 6%). vacuum freeze-drying. In the first step. Removal of alcohol salt and other low molecular weight substances can be done most efficiently with UF in the diafiltration mode. All of the other plasma proteins. the salt concentration remains at 3% in retentate and permeate since it is freely permeable to the membrane. since the solvent does not undergo a phase change. UF is preferred for protein concentration. In this approach. but it is used in subsequent processing steps. in some cases. the proteins (or other macromolecules retained by the membrane) remain at their initial concentration while the salt concentration decreases continuously. the ionic environment for the protein remains constant.Ultrafiltration
243
proteins. including antithrombin III have been concentrated as well. Buffer solutions are used to enhance its solubility. It is commonly used for fractionation as well as desalting and is practiced both as a discontinuous and as a continuous process. ethanol and salts must be removed prior to sterilizing filtration with MF and dispensing into containers suitable for I. Diafiltration refers to the process by which UF can rapidly and efficiently remove salts and/or lower molecular weight species from larger macromolecules.000 MWCO membrane so as to contain up to 25% protein.
. Figure 3. When the ethanolic precipitates are separated from their supernates. since the product need not come into contact with an air interface. In the third step. ethanol removal and desalting.000 to 30. This has been called “constant volume molecular washing” because the salts are washed out of solution. gamma-globulin is more prone to precipitation than albumin. Diafiltration. they must be redissolved and concentrated. Continuous diafiltration is generally more efficient and preferred. Continuous diafiltration reduces the processing time required in the discontinuous process. the process stream is diluted by a factor of three to reduce the salt concentration from 3 to 1%. Thus. the enzyme is now also diluted by a factor of three.V. there is less denaturation and it is more efficient than gel permeation. but freely pass through the membrane. A major advantage of concentration by UF over conventional evaporation or lyophilization is that thesalts are not concentrated by the process. delivery. a batch of the solution to be desalted is maintained at constant volume by adding pure water (dialysate) at the same rate permeate is removed. and energy costs are reduced. Gamma-globulin is also routinely concentrated by UF. The net result is the enzyme has been concentrated six-fold while the salt concentration has been reduced three-fold. the enzyme is concentrated six-fold from 3 to 18%. (2) dilution and (3) recovery. concentration to 40% has been reported. Further. Finally.) In the second step. or thin-film evaporation.

the dialysate fluid contains the microsolute intended to replace the microsolute in the starting solution. VJV. the most common method used for microsolute removal has been batch dialysis. one buffer can be exchanged for another using diafiltration. required to
. Fractionation of two macromolecules (with a large difference in molecular weight) can be accomplished more completely using diafiltration.
Co
)
“t
where
C.Ultrafiltration
245
When microsolute exchange is desired. all solutes (salts and alcohol) are removed at the same rate independent of the size and diffusivity of the various species. the same degree of salt removal can be accomplished much more rapidly with smaller volumes of dialysate. diafiltration can be used to exchange one surfactant for another in a polymer latex to enhance stabilization. In addition. In the past. The pressure driven convective transport of solutes across the membrane is much faster than concentration driven diffusion (particularly at low salt concentrations). a simple mass balance reveals that solute removal is exponential:
(31)
In
= v. Equation 31 becomes:
where
Cd = the concentration of microsolute in the dialysate feed (constant) C = the concentration of the microsolute washed in with volume Vt
Since continuous diafiltration involves the processing of a fixed volume (V. ( c. with diafiltration. avoiding problems associated with higher concentrations like low flux or denaturation. The solution to be dialyzed was placed in seamless regenerated-cellulose tubing (“sausage bags”) and suspended in the dialysate allowing salts to diffuse across the membrane. Cf Vt
= the initial microsolute concentration in the batch = the final microsolute concentration after volume Vt of wash solution (dialysate) has passed through the batch = the volume of wash solution (dialysate) added to the batch. For example.). With diafiltration. this makes the process more predictable and controllable. also equal to the volume of ultrafiltrate removed through the membrane = the initial volume of the batch. Assuming the microsolute to be removed is freely permeable to the membrane (R = 0).
V0
When microsolutes are to be exchanged or “washed in”. In the pharmaceutical industry. The entire fractionation can be carried out at low concentrations. it is convenient to speak of the number of volume turnovers.

96 may be used to predict solute clearance for a specified membrane retention (R). When the permeable species is partially rejected by the membrane (as in fractionation).95: Turnover required to achieve a specified reduction in microsolutes (R = 0) by diafiltration.
. Equation 31 predicts that a 3 volume turnover will effect a 95% reduction in microsolutes and a 5 volume turnover will wash out more than 99% of the salts.
P
a /.95. Equation 31 is plotted in Figure 3. a higher wash volume turnover will be required to achieve the same degree of clearance.96: Turnover required to achieve a specified reduction in microsolutes (having various rejections) by diafiltration.246
Handbook of Industrial Membrane Technology
achieve a fixed reduction in microsolute content.
z[/
02040 60 60 90 X OF SOLUTE REMOVED 95
Figure 3.
Percentof soluteremoved Figure 3.
6
5. Figure 3. For example.

Three types of enzyme reactors using UF membranes are introduced here. Because the processing time required for diafiltration is a minimum at IO%. the enzyme itself is not consumed in the reaction.94). If we diafilter at the starting concentration (3% in Figure 3.01 /. which utilize UF membranes in the processing of pharmaceuticals and biologicals.000 MWCO membrane without loss of activity. concentration of the virus or its proteins is frequently necessary to obtain workable volumes for subsequent processing.1 volume turnovers to reduce the salt concentration by three-fold-see Equation 31 and Figure 3. In large scale virus production.94 is 27%. serving as a catalyst-i. we must find the macrosolute concentration at which the diafiltration processing time is minimized. For example. The first of these is the enzyme reactor. the optimum concentration (C”) at which to carry out diafiltration would be 10%.94. Enzymes facilitate many biochemical reactions. In the past. the time required for diafiltration at 3% or 18% would be greater. extracts and a host of other materials can be concentrated or purified by UF.718. we will reduce the volume to be processed.
(33)
c* =
% 7
where C.. substrate is fed into one end of the column and products are withdrawn from the other end. We turn now to new developments. Ng6’ and Cooper” have derived the equation which specifies the optimum concentration (C”) at which the diafiltration process time is minimum. Enzyme Reactors.94). it does not contaminate the product and may be used over and over again. followed by concentration of the enzyme from 10% to the final concentration (18%). In fact.I and can be safely (closed system) concentrated with a 80. if we first concentrate the enzyme to its final concentration (18% in Figure 3. is the gel concentration and e is exponential e = 2.) of the enzyme of Figure 3. To optimize the process. Since diafiltration does not change the volume of the batch. On the other hand. Because the enzyme is affixed to the stationary phase. The equa-
.Ultrafiltration
247
Often. Hormones. the minimum processing time would start with concentration of the enzyme from 3 to IO%. Often continuous-flow zonal centrifuges are used for this purpose but with a significant loss in biological activity. we will achieve high flux but the volume of dialysate required for the desired salt removal will be enormous. the enzyme was discarded and fresh enzyme was charged to the reactor with substrate. diafiltration at 10% (requiring 1. Obviously. not fully commercialized. but the flux will decrease. if the gel concentration (C. Virus Concentration.e. the economics of such a reactor can be improved dramatically if the enzyme can be recovered and reused. the rota/ concentration time would be constant regardless of whether it is carried out in one step or two steps separated by diafiltration.95). One way of doing this is to immobilize the enzyme on a column of glass beads. Most viruses are larger than 0. Thus. it must be separated from the products of the reaction. the end result requires both concentration of macrosolute and reduction of microsolute concentration as in Figure 3. Enzymes may also be immobilized on membranes in the same fashion.

97 is a schematic of a continuous enzyme membrane reactor where the enzyme is free and mobile. PM-10 membrane at 15 psi.98: reactor.
IO I RETENTATE 7 l-6 E -A FILTRATE (100% GLUCOSE) I
&A
-
5
e6 :: 8 a
7
$4g
8% feed concentration. retained by a UF membrane which is permeable to the products of the reaction. Figure 3. The disadvantage is that the enzyme is more likely to become inactivated due to the continuous pumping recirculating the enzyme through the UF unit.”
FILTRATION
Figure 3.248
Handbook of Industrial Membrane Technology
tions which describe their operation may be found with a more thorough description in Chapter 7. the enzyme must be larger in size than the products.
Flow schematic of a continuous enzyme membrane reactor (mobile
Figure 3.98 shows data of Wang” on the continuous hydrolysis of starch to glucose by glucoamylase The difference in carbohydrate concentrations in the retentate and filtrate is small because the product glucose is freely permeable across the 10.) The advantage of this configuration is that there is no steric hindrance to the proper orientation of enzyme molecules with the substrate molecules. It is. however. Only the larger dextrins entering the cell as substrate are rejected. 400~(99)(105)
_
2
0
IO
I 30
40
TIEl(HOllAS1
Figure 3. Mobile Enzyme.
Continuous hydrolysis of starch by glucoamylase in membrane
. (In this scheme.97: enzyme).000 MWCO membrane.

4% feed concentration. the carbohydrate concentration in the reactor increases with time. acts as a retentive barrier to enzyme transport into the lumen of the fiber. Immobilized Enzyme. 4OOC(99)(105)
I
t
1
I
I
20
40 60 TIME (HOURS)
SO
100
Figure 3. Thus.99: Continuous hydrolysis of starch by a-amylase in membrane reactor.100: reactor. PM-10 membrane at 15 psi. Here the enzyme is immobilized within the sponge wall of UF hollow fibers. Eventually the dextrins are hydrolyzed to glucose.100 is a schematic of a continuous immobilized enzyme membrane reactor. This means it is not necessary to immobilize the enzyme onto the membrane surface unless deactivation results. on the inner wall of the fiber.99 shows Wang’s data on the hydrolysis of starch by o-amylase using the same MWCO membrane. glutaraldehyde may be used to crosslink the enzyme or cyanogen bromide may be used to covalently bond the enzyme in place.
0. Here the difference in carbohydrate concentrations between retentate and permeate is much larger because the products of a-amylase hydrolysis are predominantly maltose and larger dextrins.
Flow schematic of continuous immobilized enzyme membrane
. The membrane skin. (If immobilization is desirable. These results suggest that the selection of membrane MWCO can control the products of the reaction. but at a slower rate.)
VI PROCESS LINE (LUMEN ~1021 I”’ SUBSTRATE OR ENZYME FEED VI PROCESS BLEED
PROCESS PUMP
Figure 3.Ultrafiltration
249
Figure 3. Figure 3.

9 gms lactose/ft* 5. particles in the substrate would plug the sponge wall.0 21.0 5.0
Footnotes 1.molecular weight cutoff of 50.1 30. To achieve higher conversions. Backflush Reactor Fiber 0iameter:l Loading Density:* Substrate: Method of Operation: 20 nil 4.0% lactose Recirculating Tims Min.8 10. Both modes of operation can achieve the desired conversion. 3.5 psi outlet in this study which gave an average driving force for permeate flow of less than 1 psi.000. 23.
2.45 45 ml1 10. 0 15 60 Conversion.0% Lactose Single Pass Conversion. When the substrate is fed to the lumen of the fibers (with all permeate ports closed). Time % Min.000 and 5. 0 20 60 90 1440 B.000 are also available. Recycle Reactor3 Fiber Diameter:1 Loading Density:* Substrate: Method of Operation: 45 mil 5.0 40.9 0 1 2 3 4
gmslactose/ft2
10. Table 3. a single pass may be required.5% Lactose Recirculating Time Min. the substrate comes in contact with the enzyme in the fiber wall and product passesinto the lumen of the fiber from which it exits the module.0 3.3 9.2 11.0
Lactose Hydrolysis Via Hollow Fiber Reactors
0. These fibers have a. x 0. Table 3.45 gms lactose/fte
Conversion. Wallerstein lactase (200. When the substrate is fed to the shell side of the hollow fiber module (“back flush mode”). Note they operated the “back flush reactor” in both recirculating and single pass modes. Fibers used in this study were 20 ml1 diameter XH50 and 45 nil diameter XMSO.0 40. x 17. Number of % Passes 0. Breslau and Kilcullenn have compared the “back flush mode” with the “recycle mode” on hydrolysis of lactose with lactase. If there is significant product inhibition.72). the product stream may be taken through valve V 6 to subsequent stages to get 4 passesas in the table.6 120 180 Conversion.10:
A. 0.000 LU/gram activity) The recycle reactor was operating at 25 psi inlet.10 shows their results. it will pass from the lumen to the shell side where it contacts the enzyme and products will recycle back to the lumen (“recycle mode”) (see Figure 3. The “recycle mode” has the advantage over the “back flush mode” in that the substrate does not have to be free of suspended matter.
.8 36. Fibers with molecular weight cutoffs of 10.250
Handbook of Industrial Membrane Technology
This reactor can be operated in two ways.7 22. In the “back flush mode”.

chymotrypsin has been attached to a soluble dextran using 2-amino-4.11: Relative Activities of Immobilized Chymotrypsin
Substrate ATEE Chymotrypsin Dextrhn . Table 3.Sdichloro-s-triazine as a coupling agent. they have attached the enzyme to a soluble high MW polymer which tended to stabilize the enzyme while reducing steric hindrance and increasing access of enzyme to substrate.CHYEKITRYPSIN DExTRAN CHYP~OTRYPSIN CHWOTRYPSIN
0
10
20
30
40
50
60
70
INCUBATION TIME (HOURS)
Figure 3.
.Chymotrypsin DEAE . Though less than the free enzyme.L"‘OSE . Dunhill et aI% have explored the immobilization of enzymes to mobile supports.97. The primary disadvantage to immobilization is the restricted access of the enzyme to the substrate. As we have seen.101 for the complex. A reactor was operated continuously for two weeks at 2O’C hydrolyzing casein with no loss in enzyme activity. enzymes immobilized on a membrane have increased stability. the dextran enzyme complex had more than six times the activity of the enzyme immobilized on solid DEAE-cellulose. This complex was retained in a reactor by a 100.000 MWCO UF membrane in a configuration similar to that shown in Figure 3.101:
Immobilized enzyme stability (chymotrypsin). Specifically.11 shows the relative activities as a result of immobilization.Cellulose .Chymotrypsin 100 81 13 Casein 100 73 7
Enzyme stability at 40% is demonstrated in Figure 3. For example.Ultrafiltration 251
Disperse Soluble Immobilized Enzyme.CEl.
EFFECTOF INCUBATION TIME ON ACTIVITY AT 40'C19 . Table 3.

the organism approaches a stationary phase at about the twelfth hour due to toxic metabolite production.76 Clostridium histolyticum is cultivated in batch and continuous fermentors.102:
M
Crude Product. In the batch system.1O3. Figure 3.252
Handbook of Industrial Membrane Technology
For those reactions where the products are larger in molecular weight than the enzyme. The membrane retains the biomass while products of the fermentation are continuously withdrawn through the membrane.
The increase in cell production due to UF is shown in Figure 3. Likewise. Mobile Cells. These are discussed more thoroughly in Chapter 7. Enzymes are usually more stable within the cell wall and the living cell can regenerate itself along with the enzyme.ml of Producl
PM. are removed continuously through the membrane. for Funher Processing
Flow schematic of continuous membrane fermentor. In addition. fresh substrate is fed continuously and the toxic products are continuously removed enabling growth to continue well past the fortieth hour. one is forced to go to an immobilized enzyme system or to this system where the dextran-enzyme complex can be made larger than the products. thereby improving cell growth and fermentor productivity. the enzyme production from this fermentor is enhanced (see Figure
.102’s shows one of the earliest concepts of a continuous membrane fermentor.
Prrrrurircd Air
Fermentation Tank
l
/ o
-
’
Culturs Conccntrrtion OWl. Many of the advantages of enzyme membrane reactors are applicable to fermentors. In the continuous fermentor. In some cases.30 Membrane
3
Figure 3. A possible disadvantage is that the enzyme activity per unit reactor volume may be less in the cells than that of a pure enzyme preparation. which inhibit cell growth. a fermentor is simply an enzyme reactor using intracellular enzymes rather than extracellular enzymes. metabolic waste products. There are two primary types of membrane fermentors. Membrane Fermentors.

20 30 40 50 FERMENTATIONTIME (HOURS)
I 60
70
Figure 3. protease (the enzyme) continued to be excreted by the growing biomass for 50 hours.
Q-cFEED p”
. 20
I
FEED STARTED
I
I
_e/--
I
I -o--__o
I
$
$ 9 “J
107 _ 5-
-0---
0
--‘CONTINUOUS h BATCH
2. enzyme activity begins to decrease due to an increasingly unfavorable environment. The ratio of enzyme excreted per unit mass of cell produced in the membrane system (70 units/mg) was almost twice that from the batch system (45 units/mg). In the membrane fermentor.
6 0.-o--E&.
”
”
if!
k )_ 5 D
HFA-300
IF
membrane 37oc(99)(101)
at
100 psi.
. the organism is more efficient. This would indicate that under the more favorable environment in the membrane fermentor.INUOUS
‘o--__.
I
60
40 50 20 30 FERMENTATION TIME (HOURS)
J 70
Comparison of batch and continuous fermentors in enzyme
.
HFA-300
membrane 37oc(99)(101)
at
100 psi.103:
Comparison of batch and continuous fermentors in cell production.
IO’ 0 Figure 3.
I
IO
I
I
I
.20
I IO
I I I .104: production.Ultrafiltration
253
3 . 104) .5 -
o.76 In the stationary phase of the batch fermentation.

(2) Since the membrane is retentive of slowly biodegradable components in the sewage. Dorr-Oliver has commercialized a process which couples a UF membrane to an activated sludge reactor. He has used this bioreactor to hydrolyze lactose to glucose and for malolactic fermentations. In addition. Essentially infinite detention time is provided for the slowly biodegradable materials making it possible to metabolize raw sewage almost completely. subsequent purification steps can be simplified because the product occurs in high concentrations with lower concentrations of serum components than is the case with conventional mouse ascites fluid or suspension culture procedures.” immobilized Cc//s. see Chapter 7. Waste Treatment Though UF can do little in removing low MW contributors to BOD and COD. clarifiers are unnecessary. For more information. monoclonal antibodies.” (1) Since the membrane retains the biomass. This also means that the concentration of mixed-liquor suspended solids in the reactor can be increased from 3.000 ppm. where proteins such as monoclonal antibodies are being produced. thereby reducing the size of the reactor. Thus. The system is “self-assimilating”. viruses and hormones may be produced continuously with dramatic increases in yield. antigens. Comminuted raw sewage is fed continuously to this reactor and mixed liquor is continuously circulated through the UF unit. with no generation of excess activated sludge. Cultures can be maintained for months at a time. Products like interferon.000 to 50. Drioli78r7g has also been successful in casting cellulose acetate and polysulfone UF membranes with a thermophilic microorganism (Caldariella aciddphi/a) in the casting solution (30 to 40%).
. All of the effluent passesthrough the membrane. the hollow fibers act like natural capillaries in carrying nutrients to the cells and in removing toxic wastes. sludge wastage from such a system occurs infrequently (perhaps bimonthly or quarterly) and in very small volumes. an activated sludge reactor replaces the fermentation tank. Drioli reports significant activity compared with a similar amount of free cells and no decrease in activity after 90 days. The addition of UF to an activated sludge reactor has a number of advantages. It is also thoughts’ that the presence of this large biomass absorbs components in the raw sewage which would normally foul the membrane-perhaps irreversibly.254
Handbook of Industrial Membrane Technology
Other continuous membrane fermentors show similar results. the production rate was 120 g/hr/Q compared to less than 8 gfhrl!? for the batch system.102. In the continuous production of ethanol from Zymomonas mobilis. ” Referring to Figure 3. The organism showed good resistance to organic solvents like acetone used in the casting solution. Cells can be immobilized in the walls of UF hollow fibers and can grow to tissue-like densities (106-10’ cells /cm’) between the fibers. the required detention time may be reduced or the treatment efficiency at the same contact time may be increased. This is about 10 times higher than densities achieved in roller bottles.

Thus. The oldest were installed in industrial plants in Connecticut in 1967. The compactness of the unit (8’ x 8’ x 20’) is an asset since space is a premium. Carbon adsorbs soluble organic constituents and slime generating grease. In 1970. Furthermore. preventing pollution of the final effluent (diffusate) until the biomass has developed to the point where it can digest these pollutants.
SUMMARY In its twenty-five year (plus) life.000 gpd plant (15. Several bio-membrane plants utilizing this process have been put into operation by Dorr-Oliver. the effluent contained less than 100 coliform bacteria/100 ml. DOD born) Conventional Activated Sludge Extended Aeration Bio-membraneProcess
255
<55 <30
<IO
TSSboml <60 <70
0
The diffusate BOD and TSS values for the biomembrane process indicate an overall BOD reduction of 95% or better and a 100% reduction of TSS. The reclaimed water is not used for drinking. ultrafiltration has been shown to be a reliable and economical process for many applications. Permanent surface treatments of mem-
. (4) The addition of powdered carbon to the biomembrane loop facilitates rapid start up and prevents fouling of the membranes with oils and greases. Carbon is left in the system to scour away foulants from the surface of the membrane and is regenerated biochemically. only for washing and flushing toilets. the BOD and COD of the effluent is considerably reduced not to mention the complete elimination of all suspended solids. In addition. start up times are reduced from days to hours.000 gallons in a 12-hour period) was supplied to the city of Colorado Springs for treatment of sanitary waste originating at the tourist station located on the top of Pikes Peak. but analysis shows the effluent is of better quality than the Colorado Springs tap water. a 30. The rapid start up (3 to 4 hours) is ideal for the short tourist season on the mountain. the activated carbon acts as a catalyst to enhance the multiplication rate of the biota population in the system. coliform bacteria density determinations over a period of six months indicated that 90% of the time. Yet to be solved are the fouling and flux decay problems encountered in the processing of milk and cheese whey.Ultrafiltration (3) By virtue of the infinite detention times for large unassimilated organics and the adsorption of small pollutants within the biomass. It has eliminated severe pollution problems while recovering valuable by-products which have sometimes paid for the plant in less than a year.

From 1974 through 1980. by the middle of 1980. 260
. Between 1981 and 1984 (4 years). It was reported by the facility engineer that the use of reverse osmosis increased the yield in the manufacturing operation to such an extent that the resultant savings paid for the reverse osmosis plant in about two weeks.000 gallon per day (GPD) system was placed in operation at Texas Instruments’ electronics manufacturing facility in Dallas. I The worldwide total reverse osmosis operating capacity by the end of 1970 was 880. this capacity had more than doubled to 390.000 GPD’ and.OOO to 524. In this application.000 GPD.000. Now. Sudak
INTRODUCTION The first large industrial application of reverse osmosis occurred in 1970 when a 100. there was an accelerated rate of growth during which time the reverse osmosis capacity increased from 62. Texas.’ Reverse osmosis continued to expand and it was reported3 that the operating capacity by the end of 1984 was 524.OOO GPD. By the end of 1976.1 shows that the rate of growth in reverse osmosis capacity was slow between 1970 and 1973. an increase of 27%.OOO GPD.500.500. Figure 4.000.000. reverse osmosis operating capacity had grown to 167.OOO GPD to 412.OOO GPD.500.4
Reverse Osmosis
Richard G. virtually every electronics plant in the United States uses reverse osmosis as pretreatment in the preparation of ultrapure water.OOO GPD or an increase of 560% in the 7year period.500. the reverse osmosis plant was used to pretreat Dallas municipal watdr which was being converted to ultrapure process water by ion exchange resins. Reverse osmosis finds applications predominantly in the following areas: Industrial
-
Used to prepare industrial process water or to Process wastes. the total operating capacity increased from 412.

0
11.
power sta-
-
ooo-
YEAR
Figure 4.1:
Product Water Use
Reverse Osmosis Product Water Usage
Capacity.
Table 4. Table 4.6 38.One million gallons per day Plant
73 502
100.0
.0 2. Used for various purposes.4 2.
drinking
water
Used to prepare process water in electric tions.Reverse Osmosis Irrigation Military Municipal Power Miscellaneous
261
-
Used to upgrade waters for agricultural Used for military purposes.5
Industrial Irrigation Military Municipal Miscellaneous Power Yuma DesalLing Total * MGD.5 0. year.
Used to upgrade waters to municipal levels.1:
Reverse osmosis capacity vs.MCD* 158 2 13 191 10 55 Percent of Total 31.0 14.
purposes.13 below shows the uses of reverse osmosis as described above.

Arizona.5 Percent of Total 13. the use of reverse osmosis to prepare industrial process waters and to treat industrial waste is in a close second place. Agricultural irrigation accounts for only a very small percentage of total capacity because current sources of agricultural water are much cheaper than water from reverse osmosis.0 5.000 mg/Q.2: Source of Reverse Osmosis Feedwater Capacity.262
Handbook of Industrial Membrane Technology
Reverse osmosis finds much of its application in the production of potable water for municipal water supplies. Reverse osmosis is a process that transforms an unusable water supply into a usable resource. Brackish water is defined for the purposes of Table 4.2 only as a water that may have a TDS from that of municipal water supplies up to 10. Within the last 10 years.0
As of the end of 1984. military use of reverse osmosis should increase in the future. military services have discovered reverse osmosis to be a process ideally suited to providing drinking water in the field. In Yuma. The military use of the reverse osmosis process was small as of the end of 1984 but the U. Consequently. Table 4. This is due to the fact that early reverse osmosis membranes were incapable of single stage seawater desalination and. Government to improve the quality of the Colorado River prior to its flowing into Mexico. In fact. there is sufficient operating capacity in a number of varied applications to warrant confidence in the process. they were limited to brackish water desalination. the desalination of brackish water accounted for 82% of capacity.5% of the worldwide capacity of the reverse osmosis process. The process not only is suitable for converting brackish and seawaters into potable supplies but it will also remove toxic and biological agents which result from germ.0 100. chemical and nuclear warfare. significant advances have been made in both the flux and rejection capability of membranes and reverse osmosis is technically able to desalt seawater in a single stage.4 Although reverse osmosis is a relatively new technology. In the recent past.23 shows the different types of feedwater being processed by reverse osmosis units. Power plants have found reverse osmosis very useful in preparing boiler feedwater.S.9 26.0 67. That plant alone accounts for 14. it has been an effective competitor to the distillation process in seawater desalination.S. It is capable of renovating a broad spectrum of feedwaters from municipal water supplies that need polishing for industrial purposes to seawater that is refined into a potable water supply. thus. reverse osmosis is now beginning to replace existing distillation capacity in the Middle East. However.MGD Sea Water Waste Water Brackish Water Total 524. a desalting plant is being constructed by the U.0 82.000 mg/!L Wastewater is from industrial or municipal sources and the TDS is variable. From a technical and economic point of view the process is ca-
. Table 4. Seawater is considered to have a nominal total dissolved solids (TDS) content of 35.

The feedwater flows across the membrane surface where product water permeates through the membrane and a predetermined amount remains behind as reject. the feedwater is introduced into the reverse osmosis section where the feedwater is pressurized and routed to the reverse osmosis elements which are in pressure vessels. It has economic viability in a large number of industrial applications. The reject is discharged to waste while the product water is routed to the posttreatment section. The first section is the pretreatment section in which the feedwater is treated to meet the requirements of reverse osmosis element manufacturers and the dictates of good engineering practice. Following pretreatment.
I
PRETREATMENT
I
PRODUCT * TO INDUSTRY
Figure 4.Reverse Osmosis
263
pable of desalting a broad range of feedwaters from municipal water supplies to seawater.
BASIC PROCESS CONSIDERATIONS An industrial reverse osmosis plant usually will consist of three separate sections which are shown in Figure 4.’
. Several transport models have been proposed and these will be discussed briefly below. The third or posttreatment section treats the product water to remove carbon dioxide and adds chemicals as required for industrial use of the product water.2. the subject of salt and water transport is still controversial. A more thorough presentation is made in a paper by Soltanieh and Gill.2:
Industrial reverse osmosis.
Although operating reverse osmosis plants are commonplace.

where F. it is unfortunate that for solutions. The water passesthrough the membrane by passing from one absorbed site to the next until it reaches the other side of the membrane barrier layer. the separation of salt from water takes place. This model proposes that salt and water are separated due to physical size differences by a membrane with a pore size that lies between the sizes of salt and water. a preferential sorption for the water causes a sorbed water layer to be formed at the skin. Consequently. due to the chemical nature of the membrane skin layer in contact with the aqueous solution. such as sodium chloride in water. The energy requirements for solvent migration are much less than the energy requirements for salt migration and. The product water flow through the membrane is defined as follows: F. and the water transports through the membrane at a more rapid rate than the salt.
. This fact would seem to rule out the sieve mechanism model. A = = = A*(AP-An)
water flux through the membrane water transport coefficient pressure differential across the membrane osmotic pressure differential across the membrane
AP
An
=
=
The flow of water through a reverse osmosis membrane is primarily dependent on the applied pressure differential and the osmotic pressure differential across the membrane. The osmotic pressure is directly dependent on the salt concentration of the process stream. It is hypothesized that. Another model proposed is the wetted surface mechanism or the water clustering concept of transport. In addition. In this concept. the sizes of the salt and water molecules are almost the same. thus. homogeneous membrane surface layer. the term “net applied pressure” has come to mean the applied pressure minus the feed osmotic pressure. it is theorized that the water film at the surface of the membrane obstructs the pores and prevents salt from entering the pores. each 100 mg/Q of dissolved solids is roughly equivalent to one psi of osmotic pressure. This layer of purified water is then forced through the capillary pores by pressure.. the product stream normally leaves the reverse osmosis pressure vessels at near atmospheric pressure so that the applied pressure differential is the feed pressure. Each component in a pressurized solution dissolves in the membrane and then diffuses through the membrane. The solution-diffusion model of transport assumes a nonporous. The flow of water and salt through the membrane is uncoupled. i.264
Handbook of Industrial Membrane Technology
The sieve mechanism is the simplest and easiest model to understand. they are independent of each other. It is generally recognized that reverse osmosis membrane material is wettable and that water tends to be absorbed on the membrane by hydrogen bonding. While most laymen prefer this concept. In this model. ks a rule of thumb. the osmotic pressure of that stream is negligible.e. Another concept of water and salt transport in reverse osmosis is the preferential sorption-capillary flow mechanism. Since the product stream usually has a very low salt content. the surface of a membrane is microporous and heterogeneous at all levels of solute separation.

The membrane was operated on a 5.
PSIG
Membrane flux and rejection vs. a practical
. Rejection is poor at lower pressures and increases rapidly until it reaches an asymptote at an applied pressure of about 150 psig. In a practical system. there was no product water flow until the applied pressure exceeded the osmotic pressure (50 psi).Reverse Osmosis
265
The flow of salt or dissolved solids across the membrane is dependent on the following equation: F. These data tend to substantiate the assertion of the solution-diffusion model that flow is uncoupled. This can be attributed to a near constant flow of salt with a rapidly increasing product water flow which results in a more dilute product or in increased rejection.000 mg/ll aqueous solution of sodium chloride at 25°C. = B*AC where F. there would be a number of imperfections in the membrane and the salt flow through these capillaries would contribute to the total salt flow. The solution-diffusion model seems to represent the performance of a reverse osmosis membrane. B AC = = = salt flux salt transport coefficient salt concentration gradient across the membrane
The salt transport is primarily dependent upon the concentration of dissolved solids on each side of the membrane. the flux increased linearly as would be predicted by the above water flux equation.3:
PRESSURE. As can be seen. After this.
APPLIED Figure 4.3 shows the salt rejection and flux of a low pressure polyamide membrane as a function of applied pressure.3 was derived in membrane test cells with near perfect small areas of membrane. Figure 4.
It is noted that the data shown in Figure 4. Therefore. applied pressure.

2 1.
Figure 4. The product flow as a function of temperature may be estimated’s by using the following equation: Q 25/ Qt = ex where Q 25 Qt e
X
= = =
=
flow at 25°C or 77°F flow at temperature T.2 2.
As a rule of thumb.71828 U [l/(T+273) -l/2981
T u
= =
Temperature in “C 2723 (for cellulose acetate membranes).4 1.4 graphically shows the ratio of water flux at 25°C to water flux at other temperatures. The water transport coefficient (A) is not a constant in that it varies with temperature. “C 2.2 1. This results in little or no change in rejection as a function of
.4 2.2
0
10
20
20
40
60
TEMPERATURE
OC
Figure 4.2 0. the product water flow with constant net applied pressure will increase about 3% for each degree centigrade increase in feedwater temperature.
2.4:
Temperature correction factor.0 0.2 1.266
Handbook of industrial Membrane Technology
salt transport model must also take into account the contribution of the membrane imperfections to salt flow. Salt flux through the membrane is also directly proportional to temperature and the ratio of salt flux to water flux is essentially constant at different temperatures.6 0.4 0.0 1.

Consequently. it is economically prudent to maximize the recovery in order to minimize power consumption and the size of the pretreatment equipment. Product Feedwater percent water concentration
100
concentration
The sodium chloride rejection differs from that of other inorganic and organic dissolved solids. While the above equations are helpful in describing the reverse osmosis process. As all of the feedwater must be pretreated and pressurized. the water and salt permeation coefficients also vary as a function of pH. and membrane manufacturers will provide information and rejection data that are available for their specific membrane.0 to (2)
Seawarer Membrane-The seawater test is similar to the brackish water test except that the feedwater is an aqueous sodium chloride solution with a concentration of 35. the solubility limits of the feed constituents. the concentration of salts increases in that stream with increased recovery. In cases of sparingly soluble salts.Reverse Osmosis
267
temperature.
The results area (membrane lated as follows:
are reported as gallons per day per square flux) and as rejection of sodium chloride. the recovery of a reverse osmosis plant is established after careful consideration of the desired product quality.3
. at 50% recovery.000 mg/R and the test pressure is 800 psig. the salt concentration in the reject is about double that of the feed and at 90% recovery. The recovery of a reverse osmosis plant is defined as a percentage of feedwater that is recovered as product water. feedwater availability and reject disposal requirements. Table 4. the solubility limits may be exceeded at a high recovery. For some of the newer composite membranes. such as calcium sulfate.
(I) Brackish Water Membrane-The flux and rejection of the memis determined when the membrane is tested on a feedwater aqueous sodium chloride solution with a concentration of mglQ at an applied pressure of 420 psig (net applied pres400 psig) with a temperature of 25°C and a feed pH of from 6. the water and salt transport coefficients seldom are used to describe membrane performance. This could result in precipitation of the salt on the membrane surface resulting in decreased flux and/or increased salt passage. Most manufacturers test reverse osmosis membrane with standard solutions as described below. For instance.000 sure = 5.
foot of membrane Rejection is calcu-
R = (I-CPKF) where R CP CF = = = Rejection. an increase in recovery will increase the average salt concentration in the feed/reject stream and this produces a product water with increased salt content. Since most of the salts remain in the reject stream.
brane of an 2. In addition. Reverse osmosis is a cross-flow membrane separation process which separates a feed stream into a product stream and a reject stream.0 for 30 minutes prior to data collection. the salt concentration in the reject is nearly 10 times that of the feed.

0 0. The effect has been termed concentration polarization. This results in a boundary layer with a salt concentration which is more concentrated than the bulk stream.2 99.2 96. For example. or one minus rejection. Reverse osmosis is a cross-flow process and. then it can be seen that the salt passage for the divalent ions is about one-fifth of the salt passage for the monovalent ions.0 Rejection.4
The abovedescribed tests are conducted at a low recovery rate to minimize the effects of concentration polarization which is described below. This thin. and this results in an increased osmotic pressure at that surface.6
98. mp. and it is defined by the following equation: p = CBICM where /3 CM C8 = = = Concentration polarization Concentration in the main stream Concentration in the boundary layer
Concentration polarization increases the salt concentration at the membrane surface.5 11. the layer next to the membrane surface is laminar. nearly all of the salt remains behind in the boundary layer next to the membrane. When water permeates through the membrane. Table 4. the fluid adjacent to the membrane moves slower than the main stream. If salt passage is defined as product concentration divided by the feed concentration. membrane tests above are conducted at less than 1% recovery and tests with spiral wound elements are conducted at recoveries from 5 to 10%.4 2.4 96.4 98. While the main stream flow may be turbulent. The rejection of the divalent ions such as calcium and sulfate is much better than the rejection of the monovalent ions such as sodium and chloride. as in any dynamic hydraulic process.8 98.268
Handbook of Industrial Membrane Technology
shows typical results for a composite membrane when tested on a multicomponent solution.3 0./l 61 150 12 19 189 162 97 693
ionic Rejections
Product. mg/l 0.7 0.9 3.2 3. X 99. The increased osmotic pressure causes a drop in water flow as shown in the following equation:
. laminar flow film is called the boundary layer.3:
IOn Calcium Sodium Potassium Bicarbonate Sulfate Chloride Nitrate TDS Feed.0
97. The salt must then diffuse across the boundary layer and back into the bulk stream.

5
20.9 19.0 98.4:
Feed TDS. fl= 1.3 11. It is assumed that the membrane will deliver 20 gallons per square foot per day at a 400 psi “net applied pressure” and have a rejection of 99% when there is no concentration polarization.0 Rejection.000
20.0 90.0 97.000.0 18. Table 4.ooo 5.4 below assumes a membrane operating on various feedwaters with a total dissolved solids (TDS) content of 2. The flow of salt also increases and this can be simply shown in the following equation: Fs = B*@Cl-C2) where Fs B P Cl c2 = = = = = salt flux salt transport coefficient concentration polarization main stream salt concentration product water salt concentration
In the imperfect membrane with a small number of pores.ooo 35.0 8 .3 84.e.8 18.0 98.3 2..0
Effects of Concentration Polarization
?.Reverse Osmosis Fw = A*(P-/3n) where Fw A P P n = = = = = the water flow the water transport coefficient applied pressure concentration polarization osmotic pressure of main stream
269
The increased salt concentration at the membrane surface will also increase the tendency of sparingly soluble salts to precipitate on the membrane. gfd B = 1.0
.0 19.0
20. the increased salt concentration at the membrane surface would also result in increased salt passage through the pores which would be directly proportional to concentration polarization.5
f3 = 1.1 8 = 1.5 19.000 mg/Q.0 19.0. S= 1.4 97.5 8 = 2. mg/l Flux. The information shown in Table 4.1 B = 1.9 98.9 99.8 17.7 99.5 8 = 2.0 % 99. i. 5.1.5 97.9 98.000 and 35.

In order to achieve optimum performance. research workers at the University of Florida demonstrated. Suitable feedwater pretreatment lessens the adverse effects of temperature. and it is more resistant to compaction than either the diacetate or triacetate. that cellulose acetate possessedunique salt and water transport properties which made it potentially attractive as a reverse osmosis desalination membrane. This is done by increasing the flow rate across the membrane surface or introducing turbulence promoters into the feed/reject stream. they were prone to compact at an operating pressure of 400 psig. The early cellulose acetate membranes were only suitable for brackish water desalination as they were not capable of operating at the high pressures required for seawater desalination (see below). The first experiments were conducted with animal membranes and it wasn’t unit1 1867 that artificial membranes were employed.
REVERSE OSMOSIS MEMBRANES Osmotic phenomena have been observed since the middle of the eighteenth century. especially when the application is the desalination of high TDS waters. in a short period of time. The recommendations as to minimum flows or maximum recoveries which are specified by the reverse osmosis element manufacturer should be followed at all times. Loeb and others at the University of California at Los Angeles developed techniques to prepare cellulose acetate membranes with an economical water flux and salt rejection at moderate driving pressures. It has been demonstrated that a blend of cellulose diacetate and cellulose triacetate provides an improved membrane in that: (1) (2) (3) it is more stable than the cellulose diacetate membranes. Concentration polarization cannot be eliminated. biological attack and compaction at temperatures in excess of 85°F. The first membranes were fabricated from cellulose diacetate which is subject to hydrolysis at either a low or high pH. with the resultant loss of flux to an impractical level. In the early 1950’s. With this development. During the 1960’s. the membrane rejection was insufficient to desalt seawater in one stage.
. most membrane manufacturers will recommend a minimum feed rate to or from their elements and a maximum recovery in order to minimize the effects of concentration polarization. Cellulose acetate also compacts (densifies) at pressures above 400 psig and temperatures below 85°F. While the cellulose triacetate membranes operated quite well on a short term basis. reverse osmosis became a practical possibility. In addition. with thick films. pH and biological organisms to the point where the membrane can be used in practical reverse osmosis plants. but it can be minimized by decreasing boundary layer thickness. it has a better flux and rejection than the cellulose diacetate. A considerable amount of research has been done to develop a membrane of cellulose triacetate as this material is more stable to extremes of temperature and pH and it will better withstand chemical and biological attack.270
Handbook of Industrial Membrane Technology
The penalty of a high concentration polarization is not as severe with a low TDS feedwater as it is with a high TDS feedwater.

2% of the thickness and the remainder being an open cell porous matrix (see Figure 4. the DuPont Company screened numerous polymers to determine the suitability of materials other than cellulose acetate for use in reverse osmosis desalination. The membrane barrier layer varies in thickness from 400 to 1. from one polymer.6).
. The membrane barrier layer is a dense thin film of another polymer that is formed or deposited in a subsequent operation on the porous support. composite membrane research was initiated.5). The use of electron microscopy in the 1960’s demonstrated that the cellulose acetate membranes consisted of a relatively thin dense layer and a thicker porous layer of the same material. A composite membrane is also asymmetric but it consists of two polymer layers which are the membrane barrier layer and the porous support layer (Figure 4. all of the cellulose acetate membranes will tolerate a limited amount of residual chlorine which allows chlorine to be used for feedwater disinfection. The porous support is formed separately.5:
Porous
Matrlx
Single component asymmetric membrane. The results of this work indicated that aromatic polyamides were the “choice as the best polymer type for use in the DuPont commercial permeators”. The membrane thickness is usually about 100 micrometers with the dense layer accounting for about 0.000 angstroms. The cellulose acetate membranes are asymmetric and fabricated from a single polymer.
During the 1960’s. Shortly after the concept of an asymmetric membrane was established.’ The company was most successful in developing an asymmetric aromatic polyamide reverse osmosis membrane in a hollow fine fiber configuration which successfully competed with cellulose acetate in the market place. The porous support has a thickness of between 75 and 100 micrometers and its porosity is due to numerous small perforations through the support.
/
Figure 4. by conventional membrane casting techniques.Reverse Osmosis
271
Incidentally.

By far the most valuable technique in the formation of membrane barrier layers is interfacial polycondensation. In this method. To a large degree. the composite membranes have been improved and they exhibit a higher flux and better re-
. (2) thin film polymerization and (3) interfacial polycondensation. In this case. present day commercially available composite membranes use a polysulfone porous support.6: Support
Mlcrone
Two component composite membrane. the monomer furfuryl alcohol is polymerized in situ by adjustment of pH and temperature. It is by this method that a number of polyamides and polyurea membrane barrier layers have been formed on polysulfone. This membrane proved to be highly susceptible to oxidizing agents and is of limited value. a polymer is formed on the porous support surface at the interface of organic and aqueous phases by reaction of specific molecules dissolved in each phase. Membrane barrier layers have been formed on porous supports in the following manner:’ (1) solution.OOOAn~strome
i
4”
s
r
P
5
c”
I76 Porous Figure 4. There has been very little progress in the last five years in improving the Performance of single polymer asymmetric membranes.272
Handbook of Industrial Membrane Technology
Membrane Polyamide Serrler Layer
I
1400-. Meanwhile.l. Elements containing these membranes are available commercially. One of the earliest composite membrane systems was a porous support formed from cellulose nitrate-cellulose acetate with a membrane barrier layer of cellulose triacetate. A solution of cellulose triacetate in chloroform was deposited on the porous support and the solvent was then evaporated leaving a thin film on the porous support. Thin film polymerization was used to prepare a polyfuran membrane barrier layer on polysulfone.
Several polymers have been used as porous supports. it was fragile and expensive.
The solution coating technique was used in the preparation of the cellulose triacetate membrane discussed above. While this membrane successfully desalted seawater.

8 99.2 97. Both asymmetric cellulose acetate and thin film composite membranes were tested on lime clarified secondary effluent. other tests conducted by membrane manufacturers show that the polyurea and polyamide membrane barrier layers exhibit an organic rejection that is clearly superior to that of cellulose acetate. It has also been reported that silica rejection by the thin film composite membranes is superior to that of cellulose acetate. The pilot plants were operated at 85% recovery and the rejections reported in Table 4.Reverse Osmosis
273
jection at lower operating pressures than is available with the single component asymmetric membranes. the cellulose acetate membranes can tolerate a limited concentration of residual chlorine while the polyamide and polyurea membrane barrier layers are subject to disintegration by residual chlorine and other oxidizing agents.8 95. The superior flux and rejection capabilities of the thin film composite membrane has been demonstrated at the municipal wastewater reclamation facility of the Orange County Water District in California.5: Constituent Na NH4 SO4 Cl NOa COD TDS Rejection Comparison TFC Polyamide Membrane(lO1 99. Consequently.8 89. as noted above. Use of the average concentration would give a higher rejection in both cases. alternate methods of water disinfection or chlorination followed by dechlorination are employed in reverse osmosis systems using these membranes. the thin film composite membranes exhibit ho compaction at these pressures and a very low compaction rate at pressuresof 1.0 93. While the cellulose acetate membranes are compacted at the moderate pressures required for brackish water desalination. While the above data indicates a marginal improvement in the rejection of chemical oxygen demand (COD).4
The thin film composite membrane exhibited superior overall rejection performance in these tests.
.8 97. Table 4.6 94. which is an indication of organic content. The polyamide and polyurea composite membranes can withstand higher temperatures and larger pH variations than the cellulose acetate based membranes and they are immune to biological attack. with ammonia and nitrate rejection showing an outstanding improvement.0 94. Reverse osmosis element manufacturers should be contacted for rejection data on specific organic compounds.7 >94.5 are the percent rejection of the constituents in the feedwater and not the rejection of the average concentration of the specific constituents in the feed/reject stream.3 93.000 psig.5
CA Membranetg 93.3 71.0 99. On the other hand.

000 psig.500 feet-and as a hollow fine fiber. (3) tubular and (4) hollow fine fiber. a typically higher fluid operating pressure is used. The asymmetric cellulose acetate was originally produced as a sheet and later as a hollow fine fiber.274
Handbook of Industrial Membrane Technology PACKAGING
MEMBRANE
Reverse osmosis membrane is produced in sheet form-up to 60 inches wide and lengths up to 1. The plate and frame configuration can achieve a packaging density of up to 150 square feet of active membrane area per cubic foot of pressure vessel. sponsored much of the development of membrane packaging configurations for reverse osmosis membrane. Basically. it becomes concentrated and eventually leaves the pressure vessel as concentrate or reject. More advanced designs utilize external tie bolts which hold together sufficiently thick end plates which contain the pressurized feed/reject stream. This fabric both supports the membrane against the operating pressure and provides a flow path
.7 is a schematic of the plate and frame approach. As the feed/reject stream passes across the membrane. It has disadvantages of complex flow patterns and high costs. The configurations which have been developed and evaluated are as follows: (I) plate and frame. the plate and frame concept has not been economically competitive and there are no remaining manufacturers of this type of equipment in the United States. The asymmetric aromatic polyamide was originally produced as a hollow fine fiber and later in sheet form. Figure 4. Spiral Wound The spiral wound membrane packaging configuration is shown in Figure 4. in reverse osmosis. but research has been and will continue to be done to produce the composite reverse osmosis membranes as a hollow fine fiber. Plate and Frame The plate and frame configuration is much like the conventional plate and frame filtration concept except that. Department of Interior. Feedwater under pressure enters the top of the pressure vessel and flows between the parallel stacks of membrane/porous plates. The Office of Saline Water.
(2) spiralwound. Product water passes through the membrane and into the porous plate to be routed to the product water collection system and then out of the pressure vessel. A stack of parallel porous plates are used to support the membrane on each side of the porous plate. U.S. the spiral wound element consists of two sheets of membrane separated by a grooved. Overall. The plate and frame packaging configuration has an advantage in that only the membrane must be replaced when a membrane becomes defective. polymer reinforced fabric material. The composite membranes with polyamide or polyurea membrane barrier layers are produced in sheet form as of the end of 1987.8. The membrane package is installed in a pressure vessel designed and fabricated to withstand operating pressures from 400 to 1.

with a plastic mesh spacer between the facing membrane surfaces. The membrane envelope is rolled up around central product water tube.7: Plate and frame packaging.000 to 5. Reverse osmosis plants with a capacity from 100. The elements have an outer wrap to contain the feed/reject stream in the mesh passageway and brine seal to insure that the feed/reject stream goes through the element and not around it.000. Packaging densities of up to 300 square feet of membrane to 1 cubic foot of pressure vessel have been attained. Spiral wound elements are available in lengths from 12 to 60 inches and diameters from 2 to 12 inches.
WI
WdlW In
a
Fresh
’ Porous PIales
Figure 4. The fourth side is attached to a product water tube which has perforations within the edge seal so that product water can be removed from the porous product water carrier material. The membrane envelope is sealed with an adhesive on three sides to prevent contamination of the product water.OOO gpd of product would normally use elements with an 8-inch diameter by a 40-inch or 60-inch length. in a spiral. The mesh spacer not only serves to separate membrane surfaces but it provides a flow path for and turbulence in the feed/reject stream of each element.Reverse Osmosis
275
for egress of the product water.
.

The disadvantages are the excessive number of tube end fittings in proportion to the active membrane area in each pressure vessel.10. In a single pressure vessel with six elements. The pressure vessel inside diameter is sized to match the outside diameter of the element brine seal. The pressure vessels are designed and fabricated to accommodate from 1 to 6 elements and operating pressures from 50 to 1. The major advantages of the tubular reverse osmosis configuration are the ability to tolerate high suspended solids concentrations in the feed and the possibility of mechanical membrane cleaning. The membrane can be bonded to the tube in which case it is cast in situ or the membrane can be loose. The element product water tubes in the pressure vessel are connected to each other with interconnectors which again are O-ring devices whose seals prevent product water contamination. Pressurized feedwater enters the tube through an end fitting which seals the membrane to the tube and prevents cross contamination of the product water. The disadvantage of the element is that a moderate amount of pretreatment is required for some feedwaters to prevent fouling of the mesh brine spacers.
. the bulkiness of the reverse osmosis plant and the high cost. The reject from the first element flows to and through the second element and the reject from this element becomes the feed to the next element.9 shows a pressure vessel with 6 elements installed. between 40 and 60% of the feedwater to the pressure vessel is recovered as product water from a brackish water feed and from 25 to 35% is recovered from a seawater feed. stainless steel and fiberglass tubes have also been successfully used. The advantages of the spiral wound elements are the high packing density and high flux which makes it one of the most cost effective elements. Figure 4. Tubes with inside diameters of % and 1 inch have been used. The tube serves as the pressure vessel and the membrane is installed inside the tube. from either end of the pressure vessel. etc. usually at near ambient pressures. Tubular The tubular reverse osmosis device is shown in Figure 4. Uniformly porous fiberglass reinforced plastic tubes have been used and nonporous but perforated copper. The reject from the last element is routed from the pressure vessel to the high pressure reject manifold. The feed water flows down the length of the tube and product water permeates through the membrane and weeps through the tubular pressure vessel into a collection basin. The product water can exit the pressure vessel. The reject flows through an end fitting and is routed to additional tubes in series or to waste. The first and sixth element product water tubes are sealed to the pressure vessel end caps by O-ring devices to prevent contamination by the high pressure feed or reject to the purified product water stream. The loose membrane is cast in sheet form and a cylindrical section is formed and placed in the tube.000 psig. Feedwater enters one end of the pressure vessel and flows through the first element in which about 10% of the feed permeates through the membrane and flows through the product water carrier material into the product water tube. Packing densities for the %-inch diameter tube are about 100 square feet per cubic foot and about 50 square feet per cubic foot for the l-inch diameter tube.Reverse Osmosis
277
Spiral wound elements are installed in a pressure vessel which is usually fabricated from fiberglass reinforced plastic.

000 square feet per cubic foot but the flux is considerably lower than with sheet membrane.11 is a schematic of a typical hollow fine fiber element. As of the end of 1987. The most readily available polymers in hollow fine fiber elements are aromatic polyamide. the adhesive and protruding fibers in the product water end are machined
.10:
Tubular
packaging. Figure 4. A continuous fiber is looped into a bundle around a central feed tube.” The ratio of outside to inside diameter exceeds two. there were no large scale tubular reverse osmosis manufacturers in the United States. into hollow fine fibers. When cured. This well-known technology was adapted to prepare polymers.
There have been a number of attempts to commercialize tubular reverse osmosis systems in industrial applications. Hollow Fine Fiber
Hollow fine fiber synthetic filaments have been prepared for use in the textile industry for a long time. The aromatic polyamide hollow fine fiber has an outside diameter of 85 micrometers for the brackish water fiber and an outside diameter of 95 micrometers for the seawater fiber. cellulose diacetate and cellulose triacetate. which are suitable for reverse osmosis desalination. Both ends of the fiber bundle are potted in an adhesive. The fibers are indeed very fine in that they approach the diameter of a single human hair. The central feed tube is sealed at the product water end and is perforated within the active membrane area. The fibers are thick walled cylinders that have the compressive strength necessary to withstand the operating pressures. Both of the fibers have an internal diameter of 42 micrometers.Reverse Osmosis
279
Practical Unit Salt Water In
Fresh Water Figure
4. This concept is capable of achieving a high packaging density of 3. although there are some in Europe and Japan.

The feed flows through perforations in the tube and then radially between the spaces in the fiber toward the outer diameter of the bundle. The fiber bundle is then installed in a pressure vessel and an O-ring in the product water end tube sheet seals against the pressure vessel wail to prevent the high pressure reject stream from contaminating the low pressure product stream. The element resembles a tube and shell heat exchanger. Normally. The porous surfaces used have been carbon. each hollow fine fiber element is capable of recovering from 40 to 75% of the feed water from brackish water and from 25 to 35% from seawater. Hollow fine fiber elements are available with nominal diameters of 4. The membrane-forming materials have been hydrous thorium. and polyacrylic acid. synthetic polyelectrolytes. natural humic and fulvic acids. One of the most successful membrane-forming materials has been a layer of polyacrylic acid deposited on top of a layer of zirconium oxide. Product water permeates the membrane.
. several plants using dynamic membranes have been installed and there has been a renewed interest for such applications as orange juice concentration. through the end cap and to waste via the high pressure reject manifolding.” The dynamic membrane is formed within minutes and the optimum performance is usually achieved within an hour.Reverse Osmosis
281
to open the fibers at that end. The reject flows from the outer diameter of the fiber bund!es. Recently. the capital and operating costs have not always been commercially competitive. zirconium oxide. enters the capillaries of the fiber bundle and flows to the tube sheet where it is discharged into the product water plenum and then into the product water manifold. The advantages of the hollow fine fiber element are the high packaging density and the elimination of membrane support materials. Use of a stainless steel tube which had been pretreated with a filter aid and coated with a layer of polyacrylic acid over a layer of zirconium oxide achieved a sodium chloride rejection of 90% with a flux of about 60 gfd. However.6 and 8 inches. ceramic and stainless steel tubes. Dynamic Membranes Dynamic membranes are formed by flowing a feed solution containing 50 to 100 mg/!? of membrane-forming material tangential to a clean porous surface at velocities from 5 to 50 feet per second under pressures from 500 to 1.
PLANT DESIGN The first step in the design of an industrial reverse osmosis plant is to determine the amount of water to be treated. pulp mill wastes. peak demand. Longitudinal feed flow is minimized by the reject end being potted. product water quality. The pressure vessels are fiberglass reinforced plastic and they are designed and fabricated to accommodate operating pressures from 250 to 1.200 psi. These membrane promise a high flux with a lower rejection and a low membrane cost. Pressurized feedwater enters the element through the central feed tube which is connected to the high pressure feed manifold. municipal wastes. The prime disadvantage is the need for an efficient feedwater pretreatment to remove suspended and colloidal solids.000 psig.

calculate the SDI by using the equation below.12 and the silt density index (SDI) is determined as follows:
Anions Carbonate Bicarbonate Sulfate Chloride Nitrate Phosphate Fluoride
(I) Measure the amount of time required for 500 ml of feedwater to
flow through a 0.45 micrometer Millipore filter (47 mm in diameter) at a pressure of 30 psig. 10 and 15 minutes. IO or 15 minutes) = Initial time in seconds required to collect the 500 ml sample
= Time in seconds required to collect the second 500 ml sample after test time Tt (normally after 15 minutes).”
.6: Cations Calcium Magnesium Sodium Potassium Iron Manganese Ammonia Feedwater Analysis Other Silica Total DissolvedSolids Total SuspendedSolids Turbidity Silt DensityIndex Temperature Range Total OrganicCarbon PH The above determinations with the exception of the silt density index should be made in accordance with Standard Methods for the Examination of Water and Wastewater. The next step is to obtain an accurate and representative analysis of the feedwater. l3 A schematic diagram of the silt density index apparatus is shown in Figure 4. Knowledge of the feedwater constituents in Table 4. (3) After completion of the test.6.6 will provide sufficient information for an experienced design engineer to successfully design a reverse osmosis plant in the majority of applications. This analysis should include determination of the concentration of the feedwater constituents shown in Table 4.
(2) Allow the feedwater to continue flowing at 30 psig applied pressure and measure the time required for 500 ml to flow through the filter after 5. SDI = 100 (1 .282
Handbook of Industrial Membrane Technology
source of feedwater and reject discharge requirements. Table 4.Ti/Tf)
‘t
where SDI Tt Ti Tf = Silt Density Index = Total elapsed test time (either 5.

.Reverse Osmosis
283
Qauge at 3Opsig)
~illip~re
(0.O turbidity unit to maintain the element warranty. Pretreatment Section The pretreatment section of a reverse osmosis plant is designed: (1) to reduce suspended solids (1 micrometer or larger) to zero and to minimize the effects of colloids in membrane fouling.0 after pretreatment. the pretreated feedwater should have an SDI of 3.46 mloron
Filter -47mm
Holder dla. However.12:
Silt density index apparatus.0 or below in order for the element warranty to be effective. spiral wound elements have been used to recover municipal wastes with an SDI in excess of 5.0 or less and. (3) to add a threshold inhibitor to the feedwater and (4) to remove oxidizing compounds in the feedwater if required. One manufacturer of spiral wound element requires that the pretreated feedwater have a turbidity of less than 1. if hollow fine fiber elements are to be used. if spiral wound elements are to be used. an SDI of 5. As a general rule.0 or less. Mlllipore
FIlterI
Graduate (Measure Rate
Cylinder With Stop Watch)
Figure 4. (2) to adjust the pH of the feedwater.
Manufacturers of hollow fine fiber elements usually require that the pretreated feedwater have an SDI of 3.

suspended solids and colloids may be negligible and the pretreatment section may consist of PH adjustment and addition of a threshold inhibitor. these compounds will precipitate on the membrane with a resultant loss of membrane performance. it is concentrated in the feed/reject stream during the reverse osmosis process and it will precipitate on the membrane decreasing the flux and rejection.
(2) The flux and rejection of some composite membranes are a function of pH and the optimum pH is between 4 and 6. There are a number of filters available that have been successfully employed and these are: diatomaceous earth. A properly designed filter in these applications should be able to reduce the suspended solids load to zero. dual media and mixed media. It is necessary to conduct on-site jar tests in order to determine which coagulant aid is effective and what is the proper dose for that coagulant aid. Lowering the pH of the feedwater to between 4 and 6 converts some of the carbonate or bicarbonate ions to carbon dioxide and this prevents carbonate precipitation. Again. turbidity and/or SDI. such as the desalination of a well water. coagulant aids such as alum. single media. on-site jar tests are required to determine the proper coagulant aid. Cellulose acetate hydrolysis reduces the useful life of the membrane by increasing the flux and reducing the rejection of the membrane. dose and settling rate. (3) Many natural waters are saturated in calcium carbonate which is highly rejected by the membrane. Consequently. it may be necessary to use conventional coagulation followed by sedimentation prior to filtration. At the other extreme is the desalination of surface waters. Some of the more sparingly soluble compounds of concern are:
. municipal wastes and industrial wastes which may require all of the steps outlined above. Some feedwaters contain compounds in addition to calcium carbonate that may become saturated in the feed/reject stream and when that stream is concentrated in the reverse osmosis process. The following techniques can be used to remove suspended solids or to mitigate the effects of colloids: (1) If the feedwater is a municipal water supply or well water with an SDI of 6. ferric chloride and/or a polyelectrolyte are added to the feed stream prior to filtration. In-line coagulation may be used to reduce colloidal membrane fouling. In this technique.0 or below it may only be necessary to filter the feed.
(2)
(3)
The feedwater pH is usually adjusted to a pH of between 4 and 6 for the following reasons: (1) It is required (usually as a condition of warranty) to minimize the rate of hydrolysis of the cellulose acetate ester. In the case of a feedwater with a high suspended solids.284
Handbook of Industrial Membrane Technology
In some applications.

Reverse Osmosis Calcium sulfate Calcium phosphate Calcium fluoride Barium sulfate Strontium sulfate Silica
285
One method of preventing precipitation is to operate the reverse osmosis unit at a recovery which will not concentrate the feed/reject stream to the compound saturation level. but it is not amenable to mechanical cleaning. Another method of preventing precipitation is to add a threshold inhibitor. and they serve to disrupt the formation of a crystalline precipitate during the residence time of the feedwater in the reverse osmosis unit. The threshold inhibitors are added at the rate of 1 to 5 mg/Q of feedwater. organophosphates or phosphonates. For instance.
. It is also more difficult to clean than the spiral wound element. the hollow fine fiber has been tested and found to foul excessively on municipal wastewater reclamation applications while the spiral wound element has been operated successfully. In doing this. This element is widely used in ultrafiltration applications where the process streams contain suspended solids. such as sodium hexametaphosphate. The tubular element is the least susceptible to fouling and the easiest to clean. but there are no mechanical means available to clean this element. The spiral wound element requires less pretreatment than the hollow fine fiber element or. then it is necessary to remove it. certain polyacrylates. Element manufacturers have developed cleaning solutions for their specific membranes and elements. The tubular element can be cleaned not only by chemical action but also by mechanical means. they broaden the solubility limits of the sparingly soluble compounds. it is less susceptible to fouling. This is usually done by adding a stoichiometric excess of sodium bisulfite or sodium thiosulfate in accordance with recommendations from the membrane manufacturer. with the great number of close packed fibers. such as residual chlorine in the feed. As noted above. valves and instruments to circulate the cleaning solution through the reverse osmosis elements. Even with the best of pretreatment schemes. The hollow fine fiber element. A cleaning system should be provided with each reverse osmosis system and the cleaning system should consist of a tank to mix the cleaning solution and a pump with associated piping. Generally. Consequently. to the feedwater. the polyamide and polyurea membranes cannot tolerate an oxidizing agent. A sponge rubber ball is pumped through the tubular element with the chemical cleaning solution to scour the membrane surface. is an effective filter in itself. The hollow fine fiber element can be cleaned with chemical cleaning solutions. Each membrane packaging configuration has a different degree of susceptibility to fouling and an ease of cleaning. Consequently. if these membranes are used and the feed has a residual chlorine content. membrane elements will foul over a long period of time and they must be cleaned. there are two types of cleaning solutionsone for removal of organic foulants and another for removal of metal hydroxides. it is the most easily fouled membrane configuration and requires the most pretreatment. stated in another manner. The spiral wound element must be cleaned with chemical solutions.

then it is appropriate to forward a copy of the pretreated feed water analysis to reverse osmosis element manufacturers to obtain a prediction of product water quality. and price and delivery.000 GPD
Figure 4. about 50% of the first pass reject is recovered as product and 50% is reject which is sent to waste. Thus. the overall recovery of the unit is 75% as product. If the system recovery were from 40 to 60%. again.000 QPO
+-A Typl. pace”cT
WATER 250. the pressure vessels would be arranged in a 4-2-l array.286
Handbook of Industrial Membrane Technology
Reverse Osmosis Section Once the pretreatment study had been completed. As can be seen. then the spiral wound element should be used. Figure 4. all of the pressure vessels would be in parallel. possible problems with sparingly soluble compounds in the feedwater. recommended type of element.13 shows a flow diagram for a reverse osmosis unit with 75% recovery on a brackish feed. When the decision as to element type is made. The reject from the first pass pressure vesselsis then routed to the second pass pressure vessels where.13:
Reverse osmosis system. allowable recovery. The choice will depend on economics (element price) and desalination characteristics (flux and rejection). The pressurized feed is then pumped to the first pass pressure vessels where about 50% of the feed .0. if the system recovery were raised to between 85 and 90%. it will be possible to decide on the type of elements to be used in the reverse osmosis unit. The pretreated feed is routed to the high pressure pump where the feed pressure is raised to between 250 and 400 psig as required for brackish water desalination. However.0 or less. If the pretreated feed SDI is more than 3.
The quality of the product water is a function of the rejection as shown in the following equation:
. then either the spiral wound or hollow fine fiber elements can be used. a normal array for a 75% recovery unit is two first pass pressure vessels feeding one second pass pressure vessel or a 2-I array. If the SDI of the pretreated feed is 3.is recovered as product and 50% is reject. total number of elements required.Xl SECOND PASS
REJECT 83.

At 90% recovery.
The concentration of the total product then becomes: 1 .R) where Cp Cf R = = = Concentration of the product Concentration of the feed Rejection
287
This equation produces accurate results for a membrane sample or a small element with a low recovery.g. This means that the concentration of the feed varies throughout the membrane system.R) Y where . e.Reverse Osmosis cp = Cf * (1 .. Thus. + (1 _ R)
cp = Cf *
.R)
1
1
m. The average product water concentration is determined by the following formula: i?$ = E”(l-R) where cp B = = Average or total product water concentration Average concentration of the feed and reject streams.13 shows. However.
The average concentration of the feed and reject streams as represented by the following equation is not truly representative: 7=~ = feed concentration + reject concentration 2
As Figure 4. there are more elements in the first pass of a 75% recovery reverse osmosis system than in the second pass and the first pass elements produce more product. Cf and R are defined above and Y = system recovery.RI Y . The product water from the first membrane elements will be less concentrated than the product water from the last elements. The product water from the practical reverse osmosis system is combined in the product water manifold and its concentration is usually represented as the average product water concentration. Saltonstal114 has derived an equation to determine the average concentration of the feed and reject streams accurately and this equation is as follows: _ (1 .R. a practical reverse osmosis system is designed to recover from 25 to 90% of the feedwater. a greater percentage of the total product water is derived from the first pass and the total product water will be nearer in concentration to the first pass average than to the second pass average. 2% or less.(1 .Y)(’ C (I .Y)(l (1 . the initial membranes will have a feed which is about IO times less concentrated than the feed to the final membranes and the quality of the product water will vary incrementally throughout the system.

such as seawater is limited to a low recovery and requires a high rejection membrane. the average osmotic pressure will increase with increased recovery. since the osmotic pressure is proportional to concentration. Figure 4. t_ .
. The average feed/reject concentration increases with pressure and.14: Product
water vs. It is possible to use a lower rejection membrane and a high recovery with a low TDS feedwater. recovery.14 shows the dependence of product concentration on these variables.: z z 3 : a a
400
300
200
100
20
I
40
I
60
I
%
60
I
100
I
RECOVERY.15 shows the increase in average osmotic pressure and the decrease in flux for a membrane which would produce 20 gfd at zero recovery operating on a feed of 5. a high TDS feedwater. On the other hand. the membrane flux will decrease with increased recovery. ol E . Figure 4.000 mg/Q at 450 psi applied pressure. the product water concentration is dependent on the feed concentration.
The membrane flux is also dependent on recovery.288
Handbook
of Industrial
Membrane
Technology
Thus.
500
_ . Figure 4. Concentration polarization is assumed to be constant at all recoveries. membrane rejection and system recovery. At constant applied pressure.

It is noted that the osmotic pressure of seawater is about 400 psig and an operating pressure of between 800 and 1. the feed and reject streams are extremely corrosive. all three process streams are corrosive. General practice is to use PVC pipe and valves for low pressure piping (less than 100 psig) and 316 stainless steel pipe and valves for the high pressure (above 100 psig) process streams. Consequently. most seawater desalination reverse osmosis units are operated at a recovery of from 25 to 35%. the reject osmotic pressure would be in excess of 700 psig and the average osmotic pressure would be about 550 psig. m2ll 400 96% PSI -120 REJECTION -
PRESSURE
16
*QO
10
-80
6
-30
0
I
I
40
O
20
60
I
SO
I
*o 1 0
RECOVERY.Reverse Osmosis
26 FEED MEMBRANE NET 20
I-
‘160 6. recovery. The reverse osmosis process is relatively simple and instrumentation requirements are minimal.e. the membrane does not reject carbon dioxide and. consequently. This would present problems in obtaining a potable product of 500 mg/Q or less. the product stream contains about the same concentration of carbon dioxide as the feed and reject streams. corrosion resistant plastic or other corrosion resistant materials. The pressure vessels are fabricated with fiberglass reinforced plastic which is corrosion resistant..
%
Figure 4. If a recovery of only 50% were used.000 psig is required to obtain optimum flux and rejection. Following is a list of the minimum recommended parameters to be measured in a reverse osmosis system:
.000. The high pressure pump wetted parts should be either 316 stainless steel. Since lowering of the feed pH usually means that the feed contains significant quantities of carbon dioxide. i. In addition.15:
Flux and average osmotic pressure vs.

63 MGD of effluent is then pumped 19 kilometers to the refinery by either of two full capacity pumps which take suction from the inlet basins. the reject stream may contain valuable materials and this stream would be sent back to the process. the reject stream may require additional treatment prior to ultimate discharge. Saudi Arabia. A decarbonator is a packed column in which product water is introduced at the top while either forced or induced air is introduced at the bottom. Chlorinated secondary effluent arrives at the pumping station. The 4. The product water from a reverse osmosis unit will have a low pH and most probably a high concentration of carbon dioxide. which is adjacent to the sewage treatment plant in Riyadh. The carbon dioxide can be removed and the pH of the product increased by use of a decarbonator. In this case. through an open channel which flows to the pump station inlet basins. In industrial plants that use reverse osmosis to treat industrial wastes.290
Handbook of Industrial Membrane Technology
Feed FlOW Temperature Pressure Conductivity PH Product Flow Conductivity Other Differential Pressure across each array
A variety of control schemes can be incorporated in the design of a reverse osmosis plant. the reject stream is routed directly to waste discharge without any additional posttreatment. secondary effluent from the Riyadh sewage treatment plant. reverse osmosis and ion exchange demineralization. the reverse osmosis unit will have provided a large volume of water that is disposable or can be reused (the product) and a smaller volume of reject which can be treated more economically. and by using lime clarification. In other applications of industrial waste treatment by reverse osmosis. Figure 4. In a well-designed decarbonator.
INDUSTRIAL
REVERSE OSMOSIS AT A REFINERY
One of the most innovative industrial uses of reverse osmosis is at the Petromin Refinery in Riyadh. The air and water flow countercurrently over and around the column packing.16 is the process flow schematic for the refinery water treatment plant. However. The refinery takes an unusable municipal wastewater. the carbon dioxide content can be reduced to about 5 mg/Q in the water effluent. this subject is beyond the scope of this manual. Posttreatment Section In most plants that use reverse osmosis in the preparation of process water. The effluent is pumped through a
. filtration. it converts that useless waste into the entire process water requirements for the refinery. The carbon dioxide is stripped from the water and exits from the decarbonator at the top in the air stream.

the colloidal particles tend to repel each other and thereby limit the potential for aggregation and the colloidal suspension is said to be stabilized. lime clarification was chosen as the method to be used. Another means of particle destabilization is the use of natural or synthetic polymers which are long chain. For some time. The negatively charged colloids are adsorbed on the active sites with the resultant coagulation or growth of the particles. it has been known that suspended solids or colloids in water possessan electrical charge which is predominantly negative. On arrival at the refinery. due to their small size. if required. Each of the surge ponds is a concrete lined. Experience has shown that effective removal of the suspended solids in pretreatment is a prerequisite to efficient reverse osmosis membrane performance. the effluent is discharged into a 2 compartment splitting box which divides the flow evenly for discharge into the two on-site surge ponds. The water flows from the rapid mix basin to the flocculation basin which has four compartments. earth basin with a capacity of 5.292
Handbook of Industrial Membrane Technology
24:inch diameter cement lined carbon steel pipeline.O. it is difficult to remove them by settling. miscellaneous refinery plant streams and backwash from the refinery water treatment plant filter are returned to the surge ponds and mixed with the secondary effluent. Lime and sodium hydroxide are added to the first compartment and these are vigorously mixed with the secondary effluent in both compartments. a pig launcher and pig receiver. it is necessary to aggregate the smaller particles into larger particles which can more easily be removed by settling and filtration. high molecular weight polyelectrolytes with many active sites. much of the bicarbonate in the water reacts with the lime and forms an insoluble calcium carbonate and the magnesium in the water reacts with hydroxyl ions to form insoluble magnesium hydroxide. The surge ponds have a static air diffusion system to prevent sewage septicity and improve homogeneity in the secondary effluent. The pipeline is equipped with vacuum relief valves and air-release safety valves at the two highest Points ln the line plus air-release safety valves at the third highest point in the line. Each compartment is further divided into three sec-
. the pH of the effluent from the rapid mix basin is raised to between 10. As a result. The rapid mix basin has two compartments in series and each compartment has a high speed mixer.19 MGD with a quality similar to the secondary effluent. Chlorine can be added to the secondary effluent either at the sewage treatment plant pump station or at the surge ponds at the refinery.3 million gallons. ferric hydroxides.8 to 11 . The combined secondary effluent and the plant return streams (5. However. and/or lime have been used to destabilize colloidal suspensions and allow suspended particles to grow in size. At the same time. Therefore. At this pH. Suspended solids in secondary effluent are primarily organic in nature and. All surface waters and municipal effluents contain suspended solids as well as dissolved solids and the presence of suspended solids dictates the need for a pretreatment section. The effluent flows from the splitting box by gravity into the surge ponds.64 MGD) are pumped from the surge ponds to the rapid mix basin in the clarifier. the stabilizing forces must be neutralized and this is usually done by adding chemicals. Trivalent aluminum. Before the SLQ ended particles can coalesce. As a result of this electrical charge. The return stream flow is estimated to be 1. The pipeline is also equipped with pigging facilities (for pipeline cleaning) which consist of 2 pigs. The refinery clarification equipment has the capability of adding any of the chemicals mentioned above.

This amounts to about 0.62 MGD of effluent from this basin has a suspended solids concentration of about 2 mg/R and a TDS of about 1. The cooling tower is used to reduce reverse osmosis feedwater temperature when required. At this point.35
400
. A static mixer is installed in the line upstream of the cooling tower.14 MGD) and the remainder (5.7
Size. the cooling tower is inoperative.35 MGD with variable quality. the clarifier effluent flows by gravity to the serpentine recarbonator basin where carbon dioxide is added to reduce the pH to between 7. Effluent from the clarifier is saturated in calcium carbonate and this would precipitate on the filter media to clog the filter which is the next step of pretreatment. The sludge is thickened to about 6% in the sludge thickener and then filtered in a plate and frame filter press to about 50% solids. Operation of the lime kiln is economically marginal.6 0. The filtered sludge may then be transported to a a landfill for disposal or sent to the on-site multiple hearth lime kiln for regeneration. the treated water is pumped to the fire water storage tanks (0. In addition.5 15. mm 0. Consequently.
Table 4.0.0 3 . The gentle agitation provided in the flocculation basin promotes contact between the calcium carbonate.1. The water flows by gravity from the cooling tower to the filter aid mix station where it is possible to add coagulant aids to assist in the removal of suspended solids during filtration. The water then flows by gravity from the flocculation basin to the sedimenor clarification basin. The water then flows by gravity to the dual media filters.6 .37 MGD of sludge which is pumped to the sludge thickener tank.7:
Filter Media Description
Height
Media Coal Sand Gravel Support
mm so0 400
Inches 31. sulfuric acid can be added to the process stream to reduce the pH of the water in the event that carbon dioxide is unavailable to the recarbonation system. For a large part of the year. magnesium hydroxide and suspended matter which results in the formation of larger particles. The total amount of water from these three sources is 0. Following lime clarification and recarbonation. The enlarged floe particles settle to the bottom of this basin and they are removed as a 1% sludge by three circular rakes into a sludge basin or hopper. Coagulant aids are not being added at the filter aid mix basin at this time. The 5.48 MGD) is pumped to the water treatment plant cooling tower.7 below. Chlorine can be added to the process water at the cooling tower influent line or at the chlorine diffuser of the filter aid basin.1. The insoluble calcium carbonate and magnesium hydroxide are converted to soluble calcium and magnesium bicarbonate in the recarbonator basin.100 mg/Q.Reverse Osmosis
293
tions and all twelve sections are equipped with a slow speed turbine mechanism.5 and 8. The overflow from the sludge thickener and the filter press filtrate are returned to the rapid mix basin for reclamation. part of the water used in the regeneration of the ion exchange demineralizers is also returned to the rapid mix basin for recovery.6 . There are four separate filter basins with filter media as described in Table 4.

The primary purpose is to mitigate the possibility of calcium carbonate deposition by conversion of bicarbonate to carbon dioxide. The purpose of acid injection is twofold.48 MGD provides 1. 0. The sulfuric acid is injected to adjust the feedwater pH to a level of between 4 and 6. The product water from the primary reverse osmosis system is transferred to the cooling tower makeup and
. This removal is chemically accomplished by the addition of sodium bisulfite. the feedwater is routed to cartridge filters which serve to mix the chemicals which have been added upstream and to insure that any particles that may have escaped the gravity filters. with subsequent loss of desalination properties. About 50% of the feed to these latter pressure vesselsis recovered as product and the reject is manifolded through the flow control valve to waste. Backwash water overflows from the filtered water reservoir to the wash water storage tank. 0. The high pressure feedwater is pumped to the reverse osmosis train where the operating pressure of each train is adjusted by the pressure control valve upstream of the spiral wound element pressure vessels. Each primary train has 180elementsin 30 pressurevesselswhichare manifolded in a20-10 array. the cartridge filters do not improve the quality of the reverse osmosis feedwater to a large degree and they are not intended as continuous duty filters.06 MGD of general use water to the refinery. It then flows through a flow control valve to the filtered water reservoir by gravity. The effluent from the cartridge filters is routed to the primary reverse osmosis feed pump wet well. Sulfuric acid and sodium hexametaphosphate (SHMP) are injected in the feedwater line upstream of the cartridge filters. The pressurized water is fed to 20 pressure vesselsin parallel where about 50% of the feed is recovered as product and the reject from these pressure vessels is manifolded and fed to 10 pressure vessels in parallel. the chlorine is removed in the pretreatment system dechlorination basin.294
Handbook of Industrial Membrane Technology
The process water flows through the filter media by gravity to remove the remaining suspended solids.02 MGD of filtered water to the lime slakers at the rapid mix basin and the remaining 4. the rejection performance of the thin film composite membrane is pH sensitive and the optimum performance is at the operating pH level. After chemical addition. However. since chlorine will depolymerize the polyurea membrane barrier layer in the spiral wound element.The elements are located in fiberglass reinforced plastic pressure vessels with six elements per pressure vessel. Chlorine has been added to the feedwater upstream of reverse osmosis pretreatment. In general.13 MGD of filter backwash water. Each train is operated independently and there is a vertical turbine high pressure pump for each train mounted on the primary feed pump wet well. The primary reverse osmosis system contains five trains of equal capacity. The filter effluent of 5.27 MGD flows by gravity to the reverse osmosis dechlorination basin. The feedwater is then transferred from the dechlorination basin to the cartridge filter feed pumping station by gravity flow and it is then pumped to the cartridge filters. The chlorine level in the influent and effluent to the dechlorination basin is continuously monitored. Coincidentally. SHMP is added to the feedwater as a threshold inhibitor to inhibit the crystalline growth of sparingly soluble salts such as calcium sulfate. such as sand or other particulate matter is removed. The elements have nominal dimensions of 40 inches in length by eight inches in diameter and they contain composite membrane with a polyurea membrane barrier layer.

The product water from the secondary reverse osmosis system contains a high concentration of carbon dioxide as a result of pH adjustment in the primary reverse osmosis system which converts bicarbonate and carbonate alkalinity to carbon dioxide. The total required capacity is 1.83 MGD of product for a total product water capacity of 4. The pressurized feedwater is pumped to the secondary trains where the operating pressure of each train is adjusted by the pressure control valve up stream of the element pressure vessels which are fabricated from fiberglass reinforced plastic. The overall design recovery of each of the secondary trains is 85%. Consequently. The reject (1. i.Reverse Osmosis
295
desalt water storage tanks or to the secondary reverse osmosis system feed pump wet well. The first stage product water concentration of 80 mg/Q is adequate for cooling tower makeup and desalter process water but it is not pure enough for moderate pressure boiler feed. As in the primary reverse osmosis system. The spiral wound elements in the secondary trains are identical to the elements in the primary stage. Although the ion exchange resin would remove the carbon dioxide.20 MGD from the secondary reverse osmosis trains is of better quality than the feed to the primary reverse osmosis system and.10 MGD in the first stage. A second stage of reverse osmosis is used at the refinery to pretreat the first stage product prior to ion exchange demineralization. it is more economical to do so in the decarbonator. The reject flow of 0. The overall recovery for each train is 75% and the design capacity of each train is 0. An improved quality feed to the ion exchangers will improve not only the quality of ion exchanger effluent but it will reduce the quantity of regenerative chemicals required by the ion exchangers.71 MGD. Each secondary train has 21 pressure vessels (126 elements) which are manifolded in a 12-6-3 array. the first stage product must be further treated with ion exchange demineralization to achieve the desired purity.e. The reject from the last three pressure vessels is manifolded through the flow control valve in each train to the dechlorination basin. There are two degasifiers (decarbonators) in the refinery water treatment plant and each of these are in parallel and normally in operation.57 MGD for a total capacity of 1.33 MGD for the second stage reverse osmosis system.12 MGD) from the first stage reverse osmosis system is sent to the on-site solar evaporation pond for disposal. Each of the trains is rated at 0.31 MGD which can be produced by two trains with one train in standby. The daily requirements for product water are 3.35 MGD so there is an installed spare capacity of about 20%. The product water requirements are 2. thus.02 MGD for the refinery cooling towers and the desalter and 1. Since the membrane is “transparent” to carbon dioxide. it passes through both the primary and secondary reverse osmosis systems into the secondary system product water. one de-
. The secondary reverse osmosis system contains three trains of equal capacity. each train is operatred independently and there is a vertical turbine high pressure pump associated with each train. it is used as feed to the primary system.. Product recovery in these pressure vessels is again about 50% and the reject is manifolded and fed to the last three pressure vessels where about 40% of the feed is recovered as product. The pressurized feedwater is fed to the first 12 pressure vessels in parallel where about 50% of the feed is recovered as product and the remaining 50% is manifolded and fed to the next six pressure vessels in parallel. The recovered product is routed to the forced air decarbonators in the ion exchange demineralization system.

suspended solids and dissolved solids. A large number of reverse osmosis systems have been installed in industrial plants to prepare industrial process water with municipal water as the feed source.02 MGD of decarbonator effluent are used for regeneration of the resins and much of this is returned to the rapid mix basins for recovery. The ion exchange demineralization is accomplished in a two-step process involving treatment with both cation and anion resins in separate process vessels. instrumentation. Two transfer pumps. The ionized solutes are not all removed to the same degree by reverse osmosis any more than ion exchange resins have the same effect on all solutes. The secondary product is fed to the top of the degasifiers where it is allowed to cascade over the degasifier packing. The effluent is then passed through a Type I. a few that are not. an anion tank. It is estimated that 0. are installed to pump the decarbonated water through the ion exchange system. Since the water is at a low pH. and the necessary valves and piping for process control. the standby train is placed in operation and the depleted train is regenerated with acid and caustic.
REVERSE OSMOSIS AND ION EXCHANGE The preceding example of a reverse osmosis industrial application at a refinery showed that the process is capable of: (1) treating a feedwater with high suspended solids and dissolved solids concentrations. The result is that cations and anions are substituted by water molecules and a high purity effluent is available for use as boiler feedwater. Air is blown through the packed section from the bottom and it rises in the packed column countercurrent to the water. can be rejected to greater than 99%. and as will be discussed later. The water is first passed through the strong acid cation exchange resin (Amberlite IR-20) to exchange the cations for the hydrogen ion. carbon dioxide is transferred from the liquid phase to the gas phase. quality
Reverse osmosis also has been used to treat municipal water supplies for industrial purposes even though these supplies are generally low in turbidity. and (3) developing a number of process streams with different requirements. such as calcium. strong base anion exchange resin (Amberlite IRA-410) where the anions are exchanged for the hydroxide ion. About 0. Divalent and multivalent ions. mag nesium. this train is removed from service.01 MGD of product water are lost to evaporation in the decarbonator. When the ion exchange capacity of one train is depleted. The result is that the carbon dioxide concentration in the water is reduced to about 5 mg/ll. A significant number of these industrial applications are to either replace ion exchange demineralization or to pretreat municipal supplies prior to ion exchange demineralization. So-
. There are two full capacity ion exchange trains in the system and each train consists of a cation tank. iron and manganese.296
Handbook of Industrial Membrane Technology
gasifier is not used as a standby. sulfate. Reverse osmosis systems now commercially available will remove 95% or more of the dissolved solids normally removed in ion exchange. both full capacity. (2) reclaiming a water that is considered by most as unusable.

which also affect its performance. the variation in the quality of the ion exchange demineralized water is reduced. potassium and chloride are normally rejected to the 95% level or better. can be made more flexible through the use of reverse osmosis. In addition. Two substances that are frequently of concern in ion exchange demineralization are silica and organics. and there is an improvement in production and product quality. It was indicated in these comparisons that reverse osmosis could not compete favorably with ion exchange at dissolved solids concentrations below 700 mg/P and that its most favorable area of use would be from about 1. However. Where small point of use polishing columns are used. reverse osmosis was compared with other methods of demineralization. Silica may be the limiting factor in the efficiency of the anionic resins.200 to 5. Preliminary demineralization by reverse osmosis will make this water suitable for subsequent demineralization by ion exchange. During the 1960’s. It is. particularly as part of the process to produce high purity water. the resins may be utilized more efficiently. obvious that the variation in solids in the finished water will also be less when breakthrough occurs. The organics are frequently present in natural waters as aromatic polycarboxylic acid derivatives known as humic and fulvic acids. Concommitant results are a major decrease in the space and equipment necessary for regenerant storage and an extension of the useful life of the resins. such as in microelectronics manufacturing. are greatly reduced or are removed by reverse osmosis pretreatment. This idea has been totally refuted because some of the most successful applications of reverse osmosis.
. the better. The amount of dissolved solids in the feed to the ion exchange beds is 5% or less than when the raw water is fed directly. Reverse osmosis will frequently produce 90% or greater reductions in total silica concentrations. Additional reliability and control can be gained by measuring the solids concentration following reverse osmosis treatment. performance should be tested on the specific water to be treated since me results can be variable and the reason for differences between waters is not yet understood. As a result. This precaution virtually eliminates shock loadings on mixed bed polishing columns. It may be possible now to utilize seawater as a source of industrial process water. and (particularly in boiler feedwater applications) the lower the concentration before ion exchange demineralization.000 mg/Q dissolved solids. It is thus apparent that such an economically important factor as plant site location. When reverse osmosis is used for preliminary demineralization. An important factor to be remembered is that in some cases water supplies unsatisfactory for processing to high purity water may be the only sources available. therefore. The net effect is to reduce the number of regenerations required of the ion exchange columns by a factor of 20 or more.Reverse Osmosis
297
dium. substances difficult to remove from the resin. have been in treating low dissolved solids water. the danger of a rapid breakthrough becomes considerably reduced. but there is little change in the environmental salt budget. A concentrate or reject is produced by reverse osmosis. owing to reduced resin attrition. Water containing 200 mg/Q dissolved solids or less has been treated at costs equal to or lower than those of ion exchange alone. which may be dependent on the availability of suitable water. This results in a significant reduction in the amount of waste regenerant solutions that must be disposed of and a material reduction in the dissolved solids that might normally be discharged to the environment.

99
.8: Ion Nickel Cower Cadmium Chromate Cyanide Zinc
Reverse Osmosis Rejection Range RejectionRange. Table 4. and
(5) treatment of the wastewater in a publicly owned treatment works after discharge into a sewage system.99 96 . plating shops have modified their practices to reduce their material and economic losses.98 90 .
(2) removal of that material from the wastewater after pollution control enforcement. Even in trace quantities. While pollution control was greeted by many with coolness. humic and fulvic acids have been responsible for impairing the life of anionic resins and affecting performance of ion exchange columns. Loss of raw materials in this industry” can result in five distinct costs: (1) replacement of the plated material that was discharged to waste prior to initiation of pollution control practices.298
Handbook of industrial Membrane Technology
Where the silica concentrations in the raw water are high. Reverse osmosis is particularly suited to waste treatment in the plating industry because most of the toxic ions in the plating solutions are well rejected by commercially available membranes. (4) replacing of the process water (at times quite expensive) lost in wastewater. the resultant cost savings have shown that pollution control is not as onerous as expetted . The plating industry was one of the first industries to experience this enforced awareness. (3) disposal of the residue from item 2.95 98 . These organics are readily removed by reverse osmosis membranes.16
Table 4. In response to the plating industry’s increased awareness of the above costs.99 98 . reverse osmosis has been most effective. X 98 .98 90 . Reverse osmosis has been installed in many plating shops as a way to resource recovery and minimizing the size of waste treatment equipment and volume.
REVERSE OSMOSIS AND POLLUTION
CONTROL
Pollution control legislation has made industry aware of the economic penalty for inefficient use of raw materials.8 shows the rejection ranges of some of the more common toxic materials in the plating industry.

9.17:
Reverse osmosis in nickel plating.
REVERSE OSMOSIS AND SEAWATER
DESALINATION
It has been estimated that the oceans cover about 70% of the earth’s area and contain about 80% of the water on or in the earth. About 100 gallons/hour (GPH) are pumped from the first rinse tank through a cartridge filter and into a reverse osmosis unit. The work piece drags out plating bath to the first rinse tank. The reject stream contains 99% (59. The reverse osmosis product stream is combined with 5 GPH of tap water makeup. first rinse tank solution to the second rinse tank and second rinse tank solution to the third rinse tank. 333 mg/ll and 37 mg/ll. which is added to compensate for surface evaporation in the plating tank.400 mg/Q) of the nickel in the feed stream with 1% (32 mg/Q) remaining in the product stream. There are three rinse tanks in series and rinse water flows countercurrent to the workpiece.000 mg/Q. respectively. second and third rinse tanks have concentrations of 3. and the combined stream is returned to second rinse tank.000 mg/!. Consequently.
WORKPIECE
Figure 4. Figure 4.
. The waste stream (10 GPH) is sent to waste treatment which is a precipitation process.000 mg/ll in the Arabian Gulf.17 shows a schematic of this industrial application. The reject stream is routed through an activated carbon column to the plating bath.000 mg/lZ with the major ions being shown in Table 4. the first.? of total dissolved solids up to over 50. Seawater concentration varies from 25.000 mg/Q to the rinse tanks. The workpiece travels from the plating bath with a concentration of 270. The average composition of seawater is about 35.Reverse Osmosis
299
The nickel plating industry is a typical candidate for the use of reverse osmosis in pollution control.

with a concentration of 35.300
Handbook
of Industrial Table 4. The ferric hydroxide acts as a coagulant aid. it is to a large degree unusable and it is the constant dream of the desalination industry to desalt seawater in an economical manner that would allow the product to be used for agricultural purposes. If the reverse osmosis seawater plant were required to deliver a product water of potable quality (500 mg/Q or less). This capability has placed reverse osmosis in the seawater desalination business and reverse osmosis is beginning to replace distillation as the result of the clear cut economic advantage of reverse osmosis. The chlorine disinfects the seawater and oxidizes the ferrous ion to the ferric ion which forms insoluble ferric hydroxide. then the membrane element would require a minimum rejection of 98. has an osmotic pressure of almost 400 psig and this mandates a membrane element with high pressure capabilities. An example of reverse osmosis seawater desalination for industrial purposes is the system installed in a thermoelectric power plant in Venezuela in 1980. Spiral wound and hollow fine fiber elements have been developed with the capability of operating at 800 to 1. Ferrous chloride and chlorine are added to the filter influent line.000 mg/P.000 GPO of boiler feedwater and potable water. Seawater. A large Percentage of the coagulated particles are removed in the roughing filter which reduces the feedwater Sol from 15 to about 3.” The original segment of the plant is designed to produce 800. At higher recoveries.18. On the other hand.6% to attain that product quality at 0% recovery. There are five gravity roughing filters which have a dual media of sand and anthracite.500 1. Four of the filters are in operation while one iS
.000 2.570
While seawater is abundant.000 psig and sodium chloride rejections up to 99. Seawater is pumped from an intake channel to the roughing filters which are the first part of the pretreatment process.5%.700 142 65 33
Constituent Sodium Magnesium Calcium Potassium Chloride Sulfate Bicarbonate Bromide Other Ions Total
34. A process flow diagram for this system is shown in Figure 4. mg/l 10.9:
Membrane Technology Major Constituents of Seawater
Concentration.350 400 380 19. both pressure and rejection must increase to obtain potable quality. distillation and now reverse osmosis are desalting seawater efficiently enough to be used for preparing potable waters in affluent areas. This goal is a long way from reality.

is pumped The from filter
filters
is back-
The effluent polishing erating filters polishing
to the pressurized is anthracite filters to re-
further
reduce the SDI to less than 3.
stage is routed
siderable amount
of energy is wasted.000mg11 kgpd
4gOmgll
SECOND REVERSE 85%
STAGE I OSMOStS c. : 0 a. Each train consists of a The with a of the this storage the first with a high pressure pump feeding two subunits of 25 DuPont each plus the necessary piping. 800 kgpd I ti 3.
.18:
Power plant process water from seawater.5 is added which effluent the polishing effluent be harmful water to the polyamide to the primary reverse osto adjust the reverse os-
mosis membrane. pH to about mosis system. Sodium move residual bisulfite chlorine 6. 867 kgpd HOLDING TANK
Potable 133 490
Water kggd mgll
RECOVERY
4gOmgtt
Figure 4.000 GPD each of product water.
to be placed from which filters
in operation the roughing
when one of the operating filter clearwell filter.Reverse Osmosis
301
in standby washed. valves and instrumentation. reverse osmosis of 900 to be 37.000-gallon plant. The total water dissolved solids content
stage system
feed was reported stage was 490 tank is the which
mglR and the product
concentration
from
product
is sent to a 262.
From
Seawater
Intake
I
FILTERS
I
FIRST REVERSE 30%
STAGE OSMOSIS RECOVERY .407 37. psig.000 This units are operated at 30% recovery pressure mg/Q. There are four opmedia
and one standby to the would
and green sand. The reject from
provides
feed to the two second stage reverse osmosis trains and it water for the power A high pressure recovery turbine is not used and a con-
source of potable to waste.
Sulfuric
acid is then added to the filter is routed
and the pretreated
The first stage reverse osmosis systems consist of four trains which are capable of producing 5-g cartridge B-10 first pump filter permeators discharge 200.

Each train has 22 permeators in a 12-7-3 array and it is operated at 85% recovery. The following list enumerates a wide variety of applications for which the reverse osmosis process may be considered: (1) Municipal Potable Water General quality improvement of present supplies Upgrade total municipal supply Potable water from degraded supplies Brackish water desalination Seawater desalination Removal of nitrates. The product from the second stage has a TDS of 28 mg/ll and this is sent to the ion exchange system. heavy metals. piping. DuPont 8-9 permeators. Each train consists of a 5-p cartridge filter. About 85% of the first stage reverse osmosis system product water is pumped to the second stage reverse osmosis system which consists of two trains. fluorides. it is economical to further treat the first stage reverse osmosis product in a second stage reverse osmosis system to further decrease the load on the ion exchange resins. In this case. The reject from the second stage is of much better quality than the incoming seawater and this is routed to the roughing filter clearwell. the operating cycle would be short and chemical regenerant cost high. etc. Bottled water production Industrial Water Provide usable water where none available Brackish water desalination Seawater desalination Pure water production Industrial rinse waters Food industry Electroplating Power plant boiler feed Beverage production Medical Ultrapure water production Pharmaceutical
(2)
. The reverse osmosis system has operated well and the capacity of the plant has been increased to twice that of the initial segment described above. Preparation of boiler feedwater is done in an ion exchange system and. valves and instrumentation.
GENERAL
APPLICATIONS
OF REVERSE OSMOSIS
The above applications were specific examples that were chosen to demonstrate the versatility of the reverse osmosis process. a high pressure pump.302
Handbook of Industrial Membrane Technology
The first stage product quality meets potable standards but it is far from the quality required for boiler feedwater. with an influent of 490 mg/Q.

25 per gallon per day of product water installed. The estimated capital cost of the plant is $1. The brackish water system costs are shown in Table 4.
. a building for the reverse osmosis systems and office. a reverse osmosis system with pretreatment. The above installed capacity cost does not include the cost of land nor an independent power source. This cost includes the cost of wells. operating and maintenance costs for both a brackish water system and a seawater system.Reverse Osmosis Electronics Medical (3) Municipal Wastes Reclaim municipal wastewaters (sewage) for Ground water recharging Agricultural or landscape irrigation Industrial process water Improve effluent quality to meet discharge requirements (4) industrial Wastes Reclaim industrial wastewaters for Reuse within industrial plant Zero discharge Agricultural or landscape irrigation Removal of toxic substances prior to discharge Resource recovery Separate or concentrate valuable materials Electroplating industry Dairy industry
303
(5) Miscellaneous Production of pure water for high value crops Recovery of agricultural irrigation drainage Production of shipboard drinking water
COSTS OF REVERSE OSMOSIS There is no such thing as the typical cost of a reverse osmosis system or of the product water from that system as these costs depend on a number of things: Economic conditions Market conditions Plant size Local labor rates Chemical costs Power costs Site preparation Product quality requirements
Feedwater availability Accounting procedures
The DuPont Permasep Engineering Manual” has published a “guide” for the capital. They are based on a large brackish water system built in the southern United States in 1982.10.

36 0.10:
Total
Cost/l.18 where the product water from the first stage high pressure seawater system is treated in a second stage lower pressure brackish water system.48 $1.11:
Seawater
RO System Cost/l.304
Handbook
of Industrial
Membrane Technology
Water Cost for Brackish Water RO
Table 4.20
Data are provided for a seawater system that will produce IO million gallons per day of product water.00
Both the brackish and seawater reverse osmosis product water costs are based on 1982 costs and they are indicative of specific plants in an assumed location in the southern United States. The system is a two stage system similar to the one shown in Figure 4.O6/KWH) Chemicals Labor Maintenance Membrane Replacement Amortization (12X/20 years) Total
$1.80 0.11. Product Energy (O.22 0.90 1 d 75 $5.12 0. The estimated 1982 cost for such a system was $45 million which includes the RO system with pretreatment and a building for the RO system. The total water cost for such a plant is shown in Table 4.000 Gals.
. controls and office.
Table 4.O6/KWH) Chemicals Labor Maintenance and Repair Membrane Replacement Amortization (12%/20 years) Total $0.Product
Energy ($O.14 0. It does not include the cost of land or an independent energy source. The product water TDS is 200 mg/g.05 0.09 0.10 0.19 0.000 Gals. The cost of energy in the seawater system assumes that the reject from the first stage high pressure reverse osmosis system is sent to an energy recovery system which reduces the overall energy requirements for the total system by 31%.

A. K.
reverse osmosis to reclaim
on a large scale and to put the beneficial
water to a number
of already
demonstrated
REFERENCES 1. tain a reverse osmosis unit to improve economic come As of arid value in providing it was estimated regions and a method elements
nor does every household
Yet. that will compete with and win over other flash distillation into separation proc-
The future esses that becoming of potable fluxes plants with
of the process rests in the research that will result in the developReverse osmosis has gone a long way toward in the production with higher The development will be incorporated desalination of membranes
of a product more water lower
can do the same thing. The low pressure memfor brackish also further as an alternate with pipelines. 8. El-Ramly. and Congdon. reuse chlorine Based
systems will not totally membranes broaden
effectively
supply schemes that proliferate will
in the western
essential to progress. it is suggested that oxidizing-agentbe developed of existing in a thin plants available to RO and. wastewater will
for the RO to use
be in industrial
waste treatment
States to be followed be forced uses. Plants Inventory. El-Ramly. source of potable
municipal water. United
States Department
of Interior
(1977). Report No.
An elusive goal has been the development on limited Finally. municipal
it is apparent
the RO process will
It appears that the next market in the United the world Eventually.. in that con-
The initial
reverse osmosis has not caused deserts to bloom. This will
not only permit capacity
to win the competition East where tential rejection While resistant This will the enormous
for new seawater capacity
the RO process to replace distillation branes which industrial are being developed wastewater The reduced
plants that are being retired applications will
of these plants exists. Water
7.F.
membrane ment
(not total systems) was about $50 million. No.
reclamation costs which compete
will result from lower pressure/higher dams and other water United States. of reclaiming that the worldwide
municipal market
1985. knowledge of worldwide becomes research future. Wangnick. C. at the same time.Reverse Osmosis
305
FUTURE
PROJECTIONS projections of 20 years ago have proven to be unrealistic the tap water. Association (1981). film which composite presently membrane. provement Association (1984). the process has been of potable water to high inand industrial for wastes..A. N. economical from seawater. Report Report Supply No. 2. countries. Im-
6.
. CF. and Congdon. rejections
than multistage
and improved
single stage seawater reverse osmosis but it will allow in the Middle poand
operating
pressures. National Water Supply Improvement 3.
it appears that this goal damage play
aware
of the environmental
waste disposal that problem. more will remain elusive for the forseeable as the world caused by indiscriminate a key role in mitigating process will by application reclaimed in other
of ion specific membranes.
the applications
duce the cost and complexity sensitive membranes. Desalting Plants inventory. N. programs. Desalting Desalting Plants Inventory. reverse osmosis
process water to industry..

In nearly all thin film composite reverse osmosis membranes. the barrier layer is not actually visible other than as a smooth surface. Alongside this is shown a photomicrograph of a fracture edge of an actual membrane of this type at 1. it offers the possibility of each individual layer being tailor-made for maximum performance. However. In this photomicrograph. Some of these membranes may be considerably 307
. The microporous sublayer can be optimized for porosity. finely microporous support structure. compression resistance and strength. Petersen and John E. Indeed. The semipermeable coating can be optimized for water flux and solute rejection characteristics. Fabrication of a thin film composite membrane is typically a more expensive route to reverse osmosis membranes because it involves a two-step process versus the one-step nature of the phase inversion film casting method.5 Thin Film Composite Reverse Osmosis Membranes
Robert J. Because its thickness is so small. This contrasts with asymmetric reverse osmosis membranes in which both the barrier layer and the porous substructure are formed in a single-step phase inversion process and are integrally bonded. and hence quite fragile. varying to as low as 200 angstroms depending on the nature of the particular reverse osmosis membrane and its method of manufacture.1 contains a schematic diagram illustrating the concept of a thin film composite reverse osmosis membrane. Cadotte
A thin film composite reverse osmosis membrane can be defined as a multilayer membrane in which an ultrathin semipermeable membrane layer is deposited on a preformed. Both layers can be optimized for chemical resistance. But this does not necessarily result in fragility. Figure 5. The term “thin film composite” has the connotation that the barrier layer is extremely thin.000 magnification. This is a common result of the thin film composite approach. the chemical composition of the surface barrier layer is radically different from the chemical composition of the microporous sublayer. it cannot be seen at the magnification shown. the barrier layer may be quite thin.

lr2 The film can be formed elsewhere. There are several potential routes to the preparation of composite reverse osmosis membranes. green and red being thicker. thicknesses of 200 to 5000 angstroms have been achieved.Thin Film Composite
Reverse Osmosis Membranes
309
more rugged and chemically resistant than the typical asymmetric cellulose acetate membrane in field service. lifting it from the water surface. The most attractive approach from a commercial standpoint. even including the double layer membrane technique. has been the formation of the semipermeable membrane layer in situ by a classic “nonstirred” interfacial reaction method. For reverse osmosis membranes. it may be more correct to refer to such membranes simply as “composite” reverse osmosis membranes.’ The free floating film is transferred to a microporous support by bringing a sheet of the support into contact with the underside of the ultrathin film.4r5 and in considerable detail in several reports on government-funded research projects. In a related application involving float-cast thin cellulosic membranes for hemodialysis. usually cyclohexanone.
CELLULOSE
ACETATE
MEMBRANES
The first composite reverse osmosis membrane to be developed and described consisted of an ultrathin film of secondary cellulose acetate deposited onto a porous Loeb-Sourirajan membrane.5 @rn (25000 angstroms) were developed and used. Migration of the solvent into the water surface (as well as some loss by air evaporation) occurs. Alternatively.3j6” Figure 5. Therefore. can be used to control the thickness of the ultrathin film.“r” Ultrathin float-cast films exhibit visible light interference colors. Or it can be formed in place by plasma polymerization techniques. It should be noted that the barrier layer in asymmetric cellulose acetate membranes is itself only about 2000 angstroms thick.‘. to a level of about 5% by weight. In float-casting. a polymer such as cellulose acetate is dissolved in an aqueous solvent. Actual
.3 The ultrathin film of cellulose acetate was fabricated by a water surface float-casting technique. however. membrane polymer solution or polymer-forming reactants can be applied in a dipcoating process. blue being about 2000 angstroms in the case of cellulose acetate. whereby the ultrathin semipermeable film is formed or deposited on the microporous sublayer. A solvent is preferred which has a slight solubility in water and a specific gravity of less than 1. Several examples of membranes made by this last approach have reached commercial status.’ Numerous patents have appeared in recent years on the fabrication of gas separation membranes by float-casting. then laminated to the microporous support.” Also. gold being thinner. it shows a pronounced tendency to spread over the water surface. These can be used as a general guide for thickness. then dried or cured in place. When a casting dope of this type is allowed to flow down an inclined plane onto a quiet water surface. Continuous addition of the casting dope with mechanical drawing off of the solidified film. This has been described to some extent in the published technical literature. as was done in the earliest work on this membrane approach. two layers can be simultaneously cast and laminated to a carrier web.0 g/cc. leaving behind a floating solidified polymer film.2 illustrates this process schematically. thicknesses of up to 2.

A microporous sheet of a cellulose acetate/cellulose nitrate blend was first coated with a thin film of polyacrylic acid. and dried.3. The polyacrylic acid would dissolve in water and wash out during subsequent usage under reverse osmosis conditions.“r’s .
.310
Handbook of Industrial Membrane Technology
Figure 5. illustrated in Figure 5. intended to protect the microporous sheet from solvent attack during overcoating with the semipermeable barrier film. The result can be achieved on a large scale by the meniscus coating approach. airdried film is obtained.‘4r’5 4n this method. This coating was temporary in nature.12r13 An alternate route to ultrathin cellulose acetate membranes exists via the Carnell-Cassidy technique. An ultrathin coating of cellulose triacetate dissolved in chloroform was meniscus-coated onto the polyacrylic acid surface.16 A commercial adaptation of this process was developed by Lonsdale and Riley. A thin. which could be released from the glass plate by immersion in water. careful rate from a dilute solution of the polymer.2: Schematic diagram of the float-casting of ultrathin cellulose acetate membranes. a glass plate is mechanically withdrawn at a slow. These membranes exhibited salt rejections of as high as 99%.
measurements have been made by interferometric methods on films deposited and air-dried on glass plates.

These sheet materials included: (a) Loeb-Sourirajan asymmetric cellulose acetate mem-
. Reasons for this appear to be threefold. composite cellulose acetate membranes were technically difficult to scale up.
During the period of 1965 to 1972. continual improvements in asymmetric cellulose acetate membrane casting technology (such as the development of swelling agents and of blend membranes) brought the performance of asymmetric membranes to full equality with composite cellulose acetate membranes.Thin Film Composite Reverse Osmosis Membranes
311
MlCAOPOROuS SUBSTRATE f DIP ROLLER
POLYMER
SOLUTION
Figure 5. especially for seawater desalination. However.3:
Schematic diagram illustrating the meniscus coating technique.
MICROPOROUS BRANES
POLYSULFONE
SUPPORTS
FOR
COMPOSITE
MEM-
Early examples of cellulose acetate composite membranes used cellulose ester sheet materials as the porous underpinnings for the float-cast films. Second. efforts on them died out completely by 1975. the best data on flux and salt rejection for cellulose acetate membranes were exhibited by the composite membranes. Third. the advent of noncellulosic composite membranes in 1972 (the NS-100 membrane) offered much more promise for high performance (salt rejection and water flux). these membranes never reached commercial viability. First.

5. Surface views of the microporous polysulfone indicate pores of 200 to 300 angstroms. These pictures were obtained using freeze fractured samples in a Hitachi Model H-600 STEM electron microscope. researchers at North Star Research Institute began a search for compression-resistant microporous substrates. when switching from cellulosic substrates to the microporous polysulfone substrate.4 shows a graph comparing the flux levels and flux stability for three membranes made at that time: (a) float-cast cellulose acetate on microporous polysulfone.312
Handbook of Industrial Membrane Technology
branes formulated to have loose. These were concluded to be far from optimum because of their susceptibility to compaction at the seawater test condition of 1. patent on ultrafiltration membranes by Michaels that issued in 1971. which has also been observed in asymmetric aromatic polyamide and cellulose acetate membranes prepared by phase inversion methods. The SEM phor:igraphs show an asymmetric cross section of graded porosity-very dense at the top surface and highly porous at the bottom surface.
1500
psi
15
-
10
polysulfone Gelman GA-10
support--composite support--composite n membrane
CA CA ’ 0
5&_
CA anisotropic
0
8
OO 5
I
10 Time
I
15 (hours)
I
20
I
25
Figure 5. optimized for composite reverse osmosis membranes.4: Effect of support films on the reverse osmosis water flux of cellulose acetate membranes. The structure of a microporous polysulfone sheet.20 Figure 5. It has since become widely useful in its own right as an ultrafiltration membrane. high flux structures and (b) mixed cellulose ester microfiltration membranes in the tightest grades.S.4x 3. The improvement in membrane fluxes was readily apparent. High magnification views of the dense top layer show the anticipated nodular structure. It was subsequently included in a U.21 Concerning reverse osmosis membranes. But the broad scope of its usefulness was never fully appreciated until later. is illustrated in Figure 5. (b) float-cast cellulose acetate on a mixed cellulose ester microfilter support and (c) a standard asymmetric cellulose acetate membrane.‘” This effort resulted in the development of microporous sheets of polycarbonate (Lexan) and polysulfone (Udel).
. Polysulfone was recognized as a major improvement in the state-of-the-art of composite membranes at that time.500 psi prevalent at that time. In the fall of 1966.
I
I
1
I
membrane annealing 20 acetyl test
thickness temperature content
2000 80°C 39. it represented a key development that later enabled rapid progress to take place in noncellulosic composite membranes.5%
R
conditions
NaCl.

(c) transition region from cellular to nodular structure near film surface. (b) backside of sheet showing cellular structure. then delaminated prior to freezefracture for SEM (note fiber tracks on backside of the sheet).
. which extends through 85% of the sheet thickness. (f) high magnification view of the surface structure showing the texture of the top surface. (e) high magnification of the extreme top surface cross section. (d) dense nodular structure at the surface.Thin Film Composite
Reverse Osmosis Membranes
313
Figure 5.5: Cross section and surface of a microporous polysulfone sheet used in composite reverse osmosis membranes: (a) total cross section of a polysulfone sheet cast on a nonwoven polyester fabric.

a polyamide. MSI-400 (Membrane Systems.) refers to the same composition. optimized for brackish water operation on Mohawk-Wellton (Yuma) agricultural drainage water.4-toluenediisocyanate (TDI) or with isophthaloyl chloride (IPC). The most effective molecular weight range for this water-soluble polymer in composite membrane fabrication has been 10.% pact on the reverse osmosis scene. invented by Cadotte.5 to 1.25 PA-100 (UOP. 40% secondary and 30% tertiary amine). azacyclopropane). It is a globular molecule having a not quite statistical distribution of amine groups (30% primary. The polyamide analog exhibits somewhat higher flux and slightly lower salt rejection than the polyurea form of the membrane. A low pressure version of PA-lOO/NS-101 membrane. the descriptors NS-1 and NS-100 (North Star Research Institute) in the technical literature refer to the TDI-based polyurea membrane.0% solution of polyethylenimine). Fluid Systems Division) and NS-101 have been used to name the corresponding IPC-based polyamide.000. was developed by UOP under the name TFC-202.22r23 consisted of a microporous polysulfone sheet coated with polyethylenimine. Inc.000 to 60.This membrane. in the second case. A microporous polysulfone sheet is saturated with a water solution of the polymeric amine (0. Figure 5. then interfacially reacted with either 2.6 illustrates the preparation of the NS-100 membrane.314
Handbook of Industrial Membrane Technology MEMBRANE
NS-100 COMPOSITE
The NS-100 membrane (initially designated as NS-1) was the first noncellulosic composite membrane to appear in the published literature and have an im. a polyurea is formed.J--cH2CH2tjCH2 E2 b=o AH CH c=o kH
-CH2CH2N--CH2CH2~-
c=o
(
=o 0 =o A 4 0 =o 4
cNHH2
tH2
c=o
Strictly speaking.26 Polyethylenimine is a product of the self-condensation of the strained-ring compound ethylenimine (aziridine. Excess solution is drained off
. In the first case. The chemistry of this membrane is as follows:
-CCH2CH2~-CH2CH2NH--CH2CH2NH2
--CH2CH2r.

500 psi. This membrane gives 70% salt rejection and 55 gfd water flux under the same test conditions as above. 1. The film is then heated in an oven at 110°C for up to 15 minutes. it would wash out later during use of the membrane. The NS-100 membrane is capable of giving salt rejections in excess of 99% in tests on salt solutions simulating seawater (18 gfd. Fang and Chian used a statistically designed set of 33 experiments to produce a membrane with
. This membrane is highly sensitive to chemical attack by hypochlorite ion and hypochlorous acid in chlorinated feedwaters. Optimization studies on the NS-100 membrane were carried out by other groups in addition to North Star Research Institute. 3. 25°C).1% tolylene diisocyanate in hexane for 30 to 60 seconds.27 This membrane was claimed to be chlorine-resistant. which is too low a temperature to effectively crosslink the amine layer. if the fully formed NS-100 membrane is dried at 75’C. Wrasidlo prepared a variation of the NS-100 membrane in which the primary amine groups of polyethylenimine were cyanoethylated before reaction with isophthaloyl chloride. If the polyethylenimine were not insolubilized such as by this treatment. The polyurea layer itself is too thin to withstand high pressures without the help of the polyamine gel sublayer as an intermediate level support.Thin Film Composite Reverse Osmosis Membranes
315
TDI HEXANE
surface of polysulfone substrate
aqueous PEI coating
ultrathin polyurea barrier layer
heat-cured PEI gel
Figure 5.6:
Steps in the preparation of the NS-100 composite membrane.5% synthetic seawater. Internal crosslinking of the polyethylenimine takes place via ammonia elimination from adjacent amino groups. The amine-impregnated film is then immersed in a solution of 0. If the polyurea interfacial reaction step is omitted.
by positioning the film vertically. but performs the crucial step of crosslinking the residual unreacted polyethylenimine. This treatment not only dries and anneals the film. the resulting film will exhibit a salt rejection of 96% or less. a crosslinked polyethylenimine semipermeable barrier film is generated. and the polyethylenimine-coated polysulfone film is heat-cured as usual. but was probably not so. This forms a very thin crosslinked polyurea zone on the surface of the wet polyethylenimine layer. Also.

Modification of fabrication conditions produced a seawater version of MSI-400 capable of generating 17 gfd and 99% salt rejection on 35.31 and their development into the PA-300 and RC-100 commercial forms by Riley and coworkers.32r33 Polyepiamine.95 gfd and 99. In commercial fabrication trials. sensitivity to chlorine was extreme. being the first of a kind. and the barrier surface was thin and brittle.*s Sudak and coworkers at Membrane Systems. is the reaction product of polyepichlorohydrin with an excess of ethylenediamine.30+0. also called polyetheramine. Inc.000 ppm sodium chloride at 600 psi. overcoating of the membrane surface with a layer of water-soluble polyvinyl alcohol was practiced in order to overcome its brittleness and susceptibility to abrasion damage during handling and spiral element fabrication.30+ 1.. was by no means optimum in chemistry and performance. Perhaps the most significant outcome of this effort to date has been the discovery of composite interfacial membranes based on “polyepiamine” by Wrasidlo. Flux was only marginally attractive for desalination. A broad-based effort in both the United States and Japan has since taken place to find other polymeric amine reactants that would contribute to better properties.30
PA-300 AND RC-100 MEMBRANES The NS-100 membrane.18% rejection tested on 5.316
Handbook of Industrial Membrane Technology
10.000 ppm sodium chloride feedwater.000 ppm sodium chloride at 800 psi and 24°C.*’ This membrane demonstrated a flux of 20 gfd and a salt rejection of 97% when operated at 250 psi on a 5. also used statistically designed experiments to develop a membrane (MSI-400) capable of low pressure operation. Chian and coworkers have also published studies on the organic solute rejections of the NS-100 membrane*’ and its potential for pesticide removal from water.4toluenediisocyanate (RC-100) are shown below:
H2N CH2CH2NH2 -CH2-CH-Ck “2 &I
c
-CH2CH-O0 C”2 AHCH$H2NH&!NH
-CH2CH-OCH2 R NHCH2CH2NHC
I
I
. Its idealized structure and its reaction products with isophthaloyl chloride (PA-3001 and 2.

of course.3 99.7 93 99 93 99
(%)*
2.3 3.000 psi. though some irreversible flux decline could still of various organics were also good. a water flux of 20 gfd at 98% salt rejection was observed.3 6.3 98.9 98.9 6.0) were observed. The end result would be a highly crosslinked structure.3 94 99.
Table
5.0 5.Thin Film Composite
Reverse Osmosis Membranes
317
Reaction would take place on both the primary and secondary nitrogen groups.6 4.000 130 37 700 366 1.530 ppm sodium chloride at 400 psi.1 occur.000 ppm sodium chloride at 1.2 5. PA-300 was also postulated to possess good chlorine resistance. That is. in a brackish water test on 5. Similarly.2
Solute
Rejection 70-75 65-70 25 99.000 1. Also.1:
Reverse
Osmosis
Performance Various
of the PA-300 Organic
(ppm)
Composite
Membrane
Toward
Solute acetaidehyde acetic Alcozyme aspartic 2-butanone butyl citric acetic dimethyl ethanol ethyl glycine phenol phenol tetrachloroethylene trichlorobenzene * Conditions: 1000 psi.2 5.8 2.9 12.5 95 90 95. it was resistant to pH 3 to 12.8 6.3 99.34 Rejections These were in sharp contrast to organic rejection data on cellulose acetate membranes.3’ Subsequent experience showed it to be equally sensitive to chlorine as NS-100.6 3.4-dichlorophenoxy-
Both PA-300 and NS-100 exhibit mild cationic behavior because of the excess unreacted amino groups present in the barrier layer. Initially. and showed far better compaction resistance than cellulose acetate. fluxes of 20 to 25 gfd and salt rejections in excess of 99. it also shared the favorable characteristics of the NS-100 membrane. it possessed the capability to operate at elevated temperatures.0 5.4% (at pH 5.7 6.0
3.3 9. yet retaining some ductility because of the flexible polyether backbone chains. In other respects.500 465 220 10. as shown in Table 5. One manner in which this is evident is their propensity to absorb anionic surfactants and lose flux
.32 In tests on 35.400 100 100 104 100
l!! 5. These data represented major decreases in operating pressure while maintaining effective permeate production rates. Initial data on the PA-300 analog were excellent.0 to 6. 25’C acetate benzoate acid acid phthalate acid (soap) acid acetonitrile
Solutes
Concentration 600 190 425 2. 2 4.

% Consequently. poly epiaminohydrin was prepared by reduction of the azide derivative of polyepiiodohydrin. Kurihara and coworkers at Toray Industries prepared several aminated derivatives of polyepichlorohydrin. Even though it was attended by a variety of start-up and operating problems. a polyamide layer incorporating both the polymeric
.2 MGD plant at Jeddah. is believed to be RC-100 (the TDI-based analog) because of its greater stability and retention of salt rejection. The PA-300 membrane was the first composite reverse osmosis membrane to be used successfully in a major seawater desalination facility-the 3.5% salt rejection operating on a simulated Yuma feedwater at 200 psi. Some of these efforts have involved polymeric amines containing only secondary amino groups to reach a goal of improved chlorine resistance. Also.38 Polyepichlorohydrin was converted to polyepiiodohydrin. Membrane manufacture accounted for only a tiny fraction of its usage. A major problem was due to the nature and supply of polyepiamine itself. mostly unrelated to the membranes. salt rejections of 99. A low pressure version of PA-300. as well as replacement membrane. designated LP-300 membrane. Best salt rejections were obtained if the polymeric amine formulation contained a substantial proportion of the monomeric amines as coreactants in the interfacial reaction. 36 it continues to operate successfully at this date.318
Handbook of Industrial Membrane Technology
thereby. The Jeddah plant was truly a pioneer installation for spiral-wound composite membranes. 3-(methylamino)hexahydroazepine. Dow Chemical ceased production of this chemical. The polyepiamine polymer was observed to be unstable in shelf storage. undergoing a continuous change in viscosity. has also been developed for brackish water applications.3s Much of the original membrane in this plant. or 3-(amino)hexahydroazepine.37
OTHER INTERFACIAL
MEMBRANES
BASED ON POLYMERIC
AMINES
Various polyamines have been synthesized and evaluated in the fabrication of the NS-100 type of membrane. Whether any of them have reached commercial status cannot be determined because of the current trend to avoid publication of the compositions of new commercial reverse osmosis membranes. In tests on 3. wastewaters containing anionic surfactants cannot be economically treated with this particular type of composite membrane.5% at fluxes of 8 to 9 gfd were characteristic. A summary of UOP developments in the area of these various composite membranes as of the end of 1980 has been published by Riley and coworkers. and sometimes turning into insoluble gel.5% sodium chloride at 800 psi and 25OC. A three-zone barrier layer was produced. then reacted with either 4-(aminomethyl)piperidine. and membrane manufacturers desiring to use it were forced to develop their own preparative methods for the polyepiamine. Some difficulty was encountered in reliably producing this membrane during its earlier history. Saudi Arabia. These various compositions are described in the patent literature. It was originally developed by Dow Chemical Company as a cationic coagulant for water treatment purposes. consisting of a heat-crosslinked polyamine gel (as in NS-1001. Eventually.26 This membrane provided up to 24 gfd and 98. then formed composite polyamide membranes by interfacial reaction with isophthaloyl chloride.

which may or may not correspond to this patented composition. Interestingly.CH2 Y H (ammonium salt form)
-
p”2J$H2!e k
\’
so2 -
tCH2~cH2*02~n
I
These polymers were then interfacially reacted with di.7% salt rejection (50 and 20 gfd flux respectively) under the above test conditions.5. 600 psi. Toray has recently announced the commercial introduction of a new. 0.39 The chemistry of this polymer synthesis is shown below.4-toluenediisocyanate gave membranes with 96. Several other systems of polymeric amine-based interfacial polyamide membranes have appeared in patents by Kawaguchi and coworkers. Dynamic chlorine tests (5 ppm.Thin Film Composite Reverse Osmosis Membranes
319
and monomeric amine reactants. The patent description shows the diallylamine polymers to be polypiperidine (sixmembered ring) derivatives. Salt rejections for the polyamide examples rarely exceeded 95% in the patent examples. and an additional surface polyamide layer comprised almost solely of the monomeric amine reactant combined with the acyl halide.41 A third patent covers the use of additives to the polymeric amine phase. Another patent describes the attachment of polyfunctional amines as side groups onto linear soluble polymers via carboxamide or sulfonamide linkages. the best salt rejections were observed for the polyurea analogs.5% NaCI. One patent describes the use of amine-terminated oligomers prepared by reaction of polyepoxides with polyfunctional amines. For instance. Kawaguchi and coworkers at Teijin have prepared a series of polymers based on poly(diallyl amine). such polymers should be chlorine-resistant. high rejection polyamide composite membrane. 25°C) of 40 to 80 hours duration appeared to uphold this inference. its copolymer with sulfur dioxide. The polyurea analogs would not be resistant to chlorine.0 to 6. Because the only reactive amine groups in these polymers were of the secondary amine type. pH 6. these additives bringing about additional crosslinking of the residual amino groups through heat-curing after interfacial formation
. formed by crosslinking the polymers with isocyanates instead of acyl halides. and various terpolymers. poly(diallyl amine) reacted with 2.9 and 99.and trifunctional aromatic acyl halides to give polyamides. but there are a number of publications that show this monomer to produce preferably polypyrrolidine (five-membered ring) structures:
Cti2=yi
ykcH2
CJJ2.

taken from the first public report on the NS-100 membrane. could be interfacially reacted with isophthaloyl chloride to give a polyamide barrier layer with salt rejections of 90 to 98% in simulated seawater tests at 1. only a wastewater product.320
Handbook of Industrial Membrane Technology
of the barrier layer. Cadotte and coworkers reported that a monomeric amin. was used in the testing of such membranes. seawater salt rejections in excess of 96% could not be produced routinely.** Only the polymeric amine polyethylenimine showed development of high rejection membranes at that time. it was thought that polymeric amine was required to achieve formation of a film that would span the pores in the surface of the microporous polysulfone sheet and resist blowout under pressure tHowever.42 Examples of such additives include esters.Dramatic changes in membrane flux and salt rejection were observed. 44 This approach was claimed to provide high organic rejections simultaneously with low salt rejections. chiorohydrins. naphthalenesulfonic acid/formaldehyde condensate.5% magnesium sulfate. Unfortunately. surfactants).46r47.45 This improved membrane formation was achieved through optimization of the interfacial reaction conditions (reactant concentrations. and brackish water fluxes were too low to be attractive.3 lists the results of this approach. and a hexane phase containing 1 . For several years. one patent describes the preparation of amphoteric polyamide barrier layers containing both free carboxylate groups and ammonium groups. imidazoamides and carbamates. piperazine.5% synthetic seawater. whereas comparative data for typical aromatic diisocyanates or diacyl halides showed high rejections for both types of solutes. Yaginuma patented interfacial membranes made by condensation of polyalicyclic diisocyanates and diacyl halides with poiyethylenimine or polyepiamine.500 psi. To increase the flux of this membrane.
INTERFACIAL MEMBRANE
POLYMERIZATION
WITH
MONOMERIC
AMINES:
NS-300
The initial studies by Cadotte on interfacially formed composite polyamide membranes indicated that monomeric amines behaved poorly in this membrane fabrication approach.43 Such membranes show high rejection levels towards sucrose (92 to 99%) while freely passing sodium chloride (15 to 25% rejection). 25°C) and up to 4 gfd and 99. This membrane exhibited up to 26 gfd and 98% seawater rejection (3. acid acceptors. 1. partial or complete substitution of isophthaloyl chloride with trimesoyl chloride was examined.2% magnesium sulfate rejection (0. Finally.5 weight percent piperazine:sodium hydroxide:dodecyl sodium sulfate. in 1976.2. Table 5. No commercially available membranes corresponding to any of the above series of patents have as yet been reported. 25°C). 200 psi.O percent weight/volume of isophthaloyl chloride. However. A typical formula for membrane fabrication consisted of an aqueous phase containing 1 :I :0. with the possible exception of the Toray patent. This is illustrated in the data listed in Table 5. Improved technique after several years of experience in interfacial membrane formation was probably also a factor.500 psi. As the trimesoyl chloride content of the acyl halide reactant was increased from 0 to 100%. seawater salt rejection dropped while
.

Use was made of the trimesoyl chloride or alternate triacyl halides in the oligomer formation step.
.8 46 97. A few examples of seawater desalination membranes were obtained. This served to limit the degree of polymerization of the oligomer.1% NaCl 0.5% M&l2 0. As such. Even so. The chemistry of this membrane is shown as follows:
Synthesis of piperazine-terminated oligomers as prepolymers for interfacial membrane formation was also examined.5 lists the best performance data obtained for some piperazine oligomer membranes interfacially reacted with isophthaloyl chloride. Table 5.5% Na2S04 0.1% MgS04 0.9
Flux (gfd) 35 31 42 41 32 32
The Donnan ion effect is also reflected in the lower rejection of sodium chloride at the higher salinity level (0. The objective of these tests was to achieve single-pass seawater desalination membranes.5% vs. the presence of free carboxylate groups was avoided. Best results were seen for piperazinecyanurate prepolymers interfacially crosslinked by isophthaloyl chloride.5% NaCl 0.5% M&SO4 * Test conditions: 200 psi. a portion of the product was insoluble in water and was filtered out during preparation of the aqueous oligomeric amine solution for the interfacial reaction step. 25°C
Salt Rejection (XI 98. The resulting amine-terminated polyamide oligomers had low solubility in the solvent system and precipitated.2-dichloroethane. and diacyl chlorides in the interfacial reaction step.322
Handbook of Industrial Membrane Technology Effect of Cation and Anion Valence on Rejection of Various Salts by a Piperazine Trimesamide Interfacial Polyamide Membrane Reverse Osmosis Test Data*
Table 5.1%). 0.500 psi. 46r48Excess piperazine was reacted with di.0 70 50 97.and triacyl chlorides in an inert solvent such as 1.4:
Solute Used in Test Loop 0. but fluxes were low in view of the operating test pressure of 1.

5%
triethylamine
34
92. 98%. 70%.0 93.
1500
psi.
25’C
The use of piperazine-terminated oligomers in composite membrane formation has been explored by several working groups. low rejection. 45%.7. sucrose. Water flux of the membrane averages about 23 gfd at 225 psi and 25°C. Even high polymers containing piperazine pendant groups have been made. magnesium sulfate. wherein this type of polyamide membrane is reported to be resistant to chlorine in chlorinated feedwaters. The membrane can be operated at temperatures to 45°C and in a pH range from 3 to 11. which is closely based on the NS-300 membrane technology. In this respect.5.9 to Prepare Acid Acceptor NaOH triethylamine NaOH triethylamine N.0 99.75 ppm active chlorine at pH 5. the chloride anion.It is of some surprise to note that an interfacial polyamide membrane formed from piperazine and acyl halides is not fully chlorine resistant.
NF-40 COMPOSITE MEMBRANE A reverse osmosis membrane is commercially available. some chlorine damage was beginning to appear within 2. it would find probable use in industrial separations where
.2 98.0 93.5:
323
Reverse Osmosis Properties of Interfacial Membranes Formed of Piperazine Oligomers and lsophthaloyl Chloride Reverse
Osmosis Test Data* (%) 99.000 hours continuous exposure to 0. 93%. Partial discrimination between monovalent and divalent cations (sodium versus calcium) has also been observed for NF-40.49r50. Typical solute rejection data for this membrane are as follows: sodium chloride. as shown in Figure 5. the sulfate anion is associated with high rejection. N’-dimethylpiperazine Flux (gfd) 13 58 14 24 45 Rejection
Reactant Piperazine trimesoyl trimesoyl cyanuric cyanuric phosphorus
used
Oligomer chloride chloride chloride chloride oxychloride
1:l
trimesoyl
chloride: chloride 3. calcium chloride. As already noted for NS-300. named NF-40 (FilmTec Corporation). ” In a long term supervised test on brackish water at the Roswell Test Facility. The interest in piperazine stems in part from the work by Credali and Parrini on asymmetric poly(piperazineamide) membranes.4
isophthaloyl * Test
conditions:
synthetic
seawater.Thin Film Composite Reverse Osmosis Membranes Table 5.

its water flux and sulfate rejections are both higher.500 hours on a chlorinated tapwater containing about one part per million average chlorine content. This is illustrated in Figures 5.
NTR-7259
COMPOSITE
MEMBRANE
A similar type of membrane. Among other applications. the temperature and pH limitations of “loose” cellulose acetate asymmetric membranes are a problem.9 for two types of salts. as shown in Figure 5.7: Long term test of NS-300 membrane on chlorinated brackish water showing onset of membrane failure at 2. An interesting behavior of this kind of membrane is its response to feedwater salinity and operating pressure.500 hours due to chlorine attack. It has been operated successfully for 4. Compared to NF-40 membrane.10. The salt whey protein and lactose can be both concentrated and desalted in a single membrane operation (requires some diafiltration). a waste stream product of cheese making in which the sodium chloride content is very high (up to 7%).8 and 5. Potential applications exist in industrial separations. named NTR-7250. and long term chlorine resistance is claimed. The changes in membrane flux as a function of feedwater salinity are not explained by osmotic pressure differences across the membrane. is being marketed by Nitto Electric Industrial Company. This type of membrane will show changes in flux and salt rejection as ionic strength of the feedwater is varied. water softening.324
Handbook of Industrial Membrane Technology
700 _I i 70
600
-60
400
300
200
Figure 5.” This membrane exhibits high rejection of sodium and magnesium sulfates (95 to 98%) and low rejection of sodium chloride. and point-of-use filtration for high purity water in semiconductor chip manufactur-
. sodium chloride and magnesium sulfate. this membrane is being used in the treatment of salt whey.

Thin Film Composite Reverse Osmosis Membranes 325
NOUXiNY
1-M
Y
.

after curing at 110°C for IO minutes. 2-inch spiral element.2 ppm C12.7-1. 1.s3 This patent application describes a membrane made by interfacial reaction between a 1. This composition. the composition described in a German patent application by Kamiyama and coworkers. pH 6.
1
I
I
200 Flux at 25’C (gpd) go 100 0 00 0 0 008 o~ooo~o 0 000 000 0
.25:0.326
Handbook of Industrial Membrane Technology
Test 120 Product Conductivi (vS/cm) .4-7. Although its composition has not been revealed. 286 psi 150-190 Conditions: 0.5% by weight polyvinyl alcohol:piperazine:sodium hydroxide. It has one potential drawback in that pH resistance is limited to pH 9 on the alkaline side. is consistent with the patent application and with the pH 9 upper limit for alkaline stability:
-CH$HC+o
CIOC 0
E-t@ 0
. 12-20°C.10: Long term performance of NTR-7250 rinated tapwater.s4 a (somewhat overly simplified) composition is given for the NTR-7250 membrane consisting of both ester and piperazineamide groups. This limits its applications where alkaline feedwaters and alkaline cleaning are involved.
I 1000
2000 Elapsed
I
3000 Time (hours)
A00
53GO
I
Figure 5. shown below.ty 8o 50
C
pS/cm tapwater. In a recent review by Ohya on current progress in reverse osmosis membranes.9% at 200 psi on a 500 ppm salt solution.0.
during exposure to chlo-
ing.0% hexane solution of trimesoyl chloride and an aqueous solution containing 0. it may consist of. or be related to. exhibited 30 gfd water flux and a magnesium sulfate rejection of 98. This membrane.25:0.6 gpm brine flow.

use of isophthaloyl chloride as a partial replacement for trimesoyl chloride had relatively little effect on flux. For aliphatic polyamines such as polyethylenimine.. 1.1% nonaqueous solution of trimesoyl chloride. the area of aromatic amines in interfacial membrane formation had been neglected because of two factors: (a) the emphasis on chlorine-resistant compositions. this ridge-and-valley structure was absent.s7 FT-30 membrane is made from one of the simplest aromatic diamines: 1. and (b) poor results that had been observed in early work on interfacial aromatic polyamides. The backside of the barrier layer. As illustrated in Figure 5. Thus. This recipe was extraordinarily simple. and in many cases degraded membrane performance by lowering salt rejection. interfacially reacted with triacyl halides. and hydrodynamic parameters rather than the surface topography of the membrane.Thin Film Composite FT-30 COMPOSITE MEMBRANE
Reverse Osmosis Membranes
327
Cadotte discovered that aromatic diamines. surfactants and acid acceptors in the aromatic diamine solution were generally not beneficial. Another feature of this type of membrane was the rough surface of the membrane on a microscopic scale. Average thickness of the polyamide barrier layer is about 2000 angstroms.11. In the nonaqueous phase. gave membranes with dramatically different reverse osmosis performance characteristics than membranes based on aliphatic diamines.ss~s6 Before that time. interfacial aromatic polyamides from trimesoyl chloride and various aromatic diamines all showed a well developed “ridge and valley” structure.0% aqueous solution of the aromatic diamine and a 0. For example. The ridgeand-valley structure does not appear to cause any increased susceptibility to membrane fouling. sometimes grainy-was observed. The extensive patent network in aromatic polyamide (aramid) technology may also have been a limiting factor. The thicker sections of this aromatic polyamide membrane are apparently about as active in water permeation as the thinner sections. which favored use of secondary aliphatic amines such as piperazine. The final chemical structure of the membrane is believed to be as follows:
CIOC
G2.12. piperazine. In contrast. this type of membrane has a barrier layer thickness approximately equivalent to the skin thickness in asymmetric reverse osmosis membranes.6_hexanediamine.. shown in Figure 5. and ran quite contrary to experience with piperazine-based membranes. surfactants and acid acceptors were almost always beneficial in the NS-300 membrane system.3benzenediamine. a flat surfacesometimes smooth. and even &a’-diaminoxylene. Deposition of foulants to the point restricting flux and increasing salt rejection has been found to be controlled by feedwater quality. contained numerous micropores or passageways for exit of permeate water from the depths of the interfacial membrane. instead. but tended to decrease salt rejection and increase susceptibility to chlorine attack. n +0 0 coc’ 2 COCI
. element design. A typical recipe for an interfacially formed aromatic polyamide composite membrane comprised a 2.

Some hydrolysis of the trimesoyl chloride takes place during membrane fabrication.
. The high silica rejection may be a function of the combination of the carboxylate ionic groups and the degree of polymer crosslinking. in agreement with the crosslinked nature indicated above. ESCA studies indicated that approximately one-sixth of the carboxyl groups are present as ionic carboxylate and five-sixths of the carboxyl groups are present as amides. (b) underside of the barrier layer (foldover zone) showing the network of micropores inside the ridge-and-valley structure.12: Topside and underside of the FT-30 composite reverse osmosis membrane: (a) topside showing well-developed ridge-and-valley structure. believed to be approximately as follows:
Both membranes show similar behavior toward oxidants such as chlorine and bromine. leading to the above structure.Thin Film Composite Reverse Osmosis Membranes
329
Figure 5. halogen oxidation and pH extremes.‘s and both membranes show approximately equivalent behavior in organic rejections. which is advantageous in ultrapure semiconductor water preparation. The FT-30 barrier layer is insoluble in sulfuric acid and in all organic solvents.s9 The crosslinking inherent in the FT-30 membrane barrier layer appears to confer greater membrane stability towards compaction. The FT-30 membrane also appears to have a uniquely high rejection (95% or higher) of soluble reactive silica. and also an area of membrane barrier layer folded over upon itself. Its chemical structure is somewhat similar to the composition of the duPont Permasep B-9 hollow fiber polyamide.

Similar flux levels are possible with the TFC202 and LP-300 membranes. manufactured under license to FilmTec.63 Versions of this membrane. fluxes were 4 to 11 gfd. Applications for FT-30 membrane have appeared in all reverse osmosis fields from seawater desalination to home tapwater systems operating on line pressure.13.14 shows the effect of ferric ion in promoting chlorine attack. The prepolymer approach which worked well for piperazineamide interfacial membranes was not useful in the case of FT-30 and its analogs. A peculiar property of the FT-30 membrane in regard to chlorine attack was that the rate of oxidation was lowest in an acid pH range of 5 to 6 and higher in the alkaline pH range.
. whether static jar storage tests or dynamic tests with chlorine added to the feedwater. A membrane is obtained by reaction of this intermediate with trimesoyl chloride. including reverse osmosis performance under seawater and brackish water test conditions.5 to 99. showed a much lower rate of oxidation compared to other polyamide membranes such as the NS-100. This effect has also been observed by Glater and coworkers. followed by a curing step at 1 IO0 to 13O’C. In brackish water applications. Salt rejections of 98.66 This prepolymer contains a free carboxylate group and is soluble in water as the sodium salt. A patent application has appeared on the use of prepolymer formed from 1. Chlorine attack on reverse osmosis membranes is believed to be catalyzed by transition metal ions such as iron.60”2 In commercially produced spiral-wound elements . De Danske Sukkerfabrikker). At this date.‘s The pH effect on chlorine oxidation of FT-30 membrane was opposite of that normally observed with other membranes.1 to 99. Patterson Candy International) and plate-and-frame design (HR-95. Figure 5. because the hypochlorous acid at pH 5 to 6 would have a higher oxidation potential than sodium hypochlorite in the alkaline pH range. But it is notable that those membranes achieve such high fluxes through use of extremely thin surface barrier layers about only one-tenth the thickness of the FT-30 barrier layer. At pH 3. Chlorine resistance tests on FT-30 membranes. using in this case FT.64t65. as mentioned earlier. the FT30 membrane typically gives 99. manganese and cobalt.3benzenediamine and trimellitic anhydride acid chloride. the effect of hypochlorous acid on FT-30 membrane lifetime could not be measured because the underlying polysulfone support layer was instead rapidly attacked and weakened.330
Handbook of Industrial Membrane Technology
The properties of FT-30 membranes have been reviewed in several publications.’ This is illustrated in Figure 5.3% salt rejection at 24 gfd flux in seawater desalination at 800 psi and 25°C. The presence or absence of these heavy metal ions may explain the discrepancies in chlorine resistance that have been reported for basically identical membranes such as NS-300 and NTR-7250. HR-98.000 ppm sodium chloride solution at 200 psi were obtained. FT-30 spiral elements can be operated at system pressures of as low as 225 psi while producing water at 22 to 24 gfd. Poor solubility of aromatic amide precursors in aqueous media was the main obstacle.1% on 2. are available in tubular form (ZF-99. it is the only commercial reverse osmosis membrane other than cellulose acetate that has specific FDA approval for food contact usage.30 membranes preimpregnated with ferric chloride.

depending on feedwater composition.W. Mass transfer rate of diamine across the interface into the organic phase was noted to be the rate-controlling step at all concentrations of diamine. Several proofs are presented in support of this statement. on the organic solvent side of the interface. if the powder is deposited on the aqueous phase side of the growing interfacial film. Continued buildup of the membrane material on the organic side becomes impossible. For instance. 85 to 90% magnesium sulfate rejection. becomes incorporated into the polyamide film if the powder is introduced on the organic phase side. particularly when relative reactivities of the amines were comparable. In the case of a polymeric amine such as polyethylenimine or polyepiamine. the solubility of the polymer in the organic phase would be very low. removal of trihalomethane precursors from groundwaters. 67 interfacial polyamide formation is stated to occur in the organic phase. it remains loose and unattached. and point-of-use filtration of high purity water for semiconductor chip manufacturing.9. The NF-50 membrane has approximately the same characteristics as NTR-7250 and NF-40. partial water softening. but possessesan extremely high water flux.332
Handbook of Industrial Membrane Technology MEMBRANE
NF-50 COMPOSITE
A new variation related to the FT-30 membrane is being developed-the NF50 composite membrane-which would appear to occupy a unique place in membrane technology. by P. Reaction would take place at the interface to form an extremely thin. The effect of increasing salinity on salt rejection and water flux is similar to the behavior observed for NF-40 membrane as was illustrated in Figures 5. The NF-50 membrane thus becomes the first example of a reverse osmosis membrane capable of operation at ultrafiltration membrane pressures. A visually graphic display of this mechanism is illustrated in Figure 5. An insoluble colored powder. crosslinked network. In this latter case. Thereafter. Reverse osmosis operation in large systems at a pressure of 35 to 50 psi is possible. As a result. This network would block the transport of further polymeric amine from the aqueous phase into the organic phase. Condensation Polymers: By Interfacial and Solution Methods.
MECHANISM
OF INTERFACIAL
MEMBRANE
FORMATION
In the book. However. that is. 98% sucrose rejection. Water flux is 15 to 25 gfd at 35 to 50 psi.
. deposited on the face of an interfacial polyamide film. monofunctional acyl halides added to the difunctional acyl halide in the organic phase always lowered polymer molecular weights. Morgan. However.8 and 5. Applications for NF-50 membrane potentially include industrial separations. Characteristics of this membrane include 30 to 40% sodium chloride rejection. partition coefficients became a factor. and 99% raffinose rejection. monofunctional amines added to the difunctional amines in the aqueous layer did not always show this effect.15. composite membranes made by the NS-lOO/NS-101 type of approach will naturally tend to have very thin barrier layers-typically 200 to 250 angstroms thick. the growth in thickness of the barrier layer is controlled by the much slower diffusion of acyl halide or isocyanate into the aqueous phase.

By contrast. on the other hand. 2. monomeric amines lead to thicker barrier layers. combinations of a polymeric amine with a monomeric amine were used to produce composite polyamide membranes having high salt rejections.‘. molecular weight. barrier layer thicknesses as high as 2500 angstroms are readily produced. so that membrane material continues to form on the organic side. will show variations in purity. a membrane such as PA-300 is normally overcoated with a protective layer of water-soluble polyvinyl alcohol to minimize abrasion and salt rejection lossesduring spiral element assembly. transport of the amine across the water-solvent interface takes place readily.4 . and 4-aminomethylpiperidine) in interfacial polyamide membranes. monomeric amines can be obtained in most cases as pure crystalline compounds. Therefore..Thin Film Composite Reverse Osmosis Membranes
333
sebacyl chloride in xyene * ‘a. and a subsurface polyamide zone incorporating both monomeric and polymeric amine.
of the direction of growth of a polyamide
In the case of monomeric amines such as piperazine or 1.H. chain branching and viscosity from lot to lot.5dimethylpiperazine. This patent also claimed very high salt rejections. which consequently tend to show better abrasion resistance and greater tolerance to chemical attack. Polymeric amines. In the patent by Kurihara. identical in lot after lot. Uemura and Okada. First.‘AL_ __--L 20 hr insoluble polyamide -interface dye powder
-
-
-
-Ic
1. This patent disclosure demonstrated an understanding of the mechanism of interfacial polyamide barrier layer formation. A patent has also appeared by Fukuchi and coworkers on the combination of polyethylenimine and several monomeric amines (including piperazine.15: Visual demonstration film at the solution interface. the interfacially formed polymer film remains rather porous to salts and small molecules (until dried or heat-cured). Furthermore. Second.6_hexanediamine in water
sebacyl carbon
chloride tetrachloride
in
Figure 5. Monomeric amines have two advantages over polymeric amines in interfacial composite membrane fabrication. The membranes were described as having a bilayer polyamide barrier film: a surface polyamide zone rich in monomeric amine. This adds an element of variability to the membrane fabrication process.3-benzenediamine.6-hexanediamine in water
membrane
1 min
1.
SULFONATED
POLYMER
COMPOSITES:
NS-200 MEMBRANE
Several composite membranes have been prepared based on sulfonated polymers. These are typically formed on microporous polysulfone supports by solu-
.

8 to 99. One of the earliest and most unique examples of this type of membrane was the NS-200 membrane. Oxidation by dissolved oxygen is another potential
.334
Handbook of Industrial Membrane Technology
tion-coating techniques. The furfuryl alcohol was fairly rapidly converted to a resinous product in the aqueous solution. turning brown in the wet state. Excess coating solution was drained away. This hydrolysis would entail some depolymerization of the resin and concurrent formation of ionic sulfonate groups.9% salt rejection. in flat sheet form. Typical performance in a 1. Examples of additives included saccharides. In subsequent efforts on the commercialization of this membrane. 7or71145 pared by dipping a microporous polysulfone sheet in a 2:2:1 solution of furfury1 alcohol:sulfuric acid:Carbowax 20M (Union Carbide) polyethylene oxide dissolved in 80:20 water:isopropanol. It is probably severely degraded by sulfuric acid during the oven treatment step. causing the membrane to absorb water and swell in monovalent salt solutions. laboratory samples of NS-200 membranes with fluxes of 40 to 50 gfd under these same test conditions were obtained.7o only furfuryl alcohol and sulfuric acid in water were used. In simulated seawater tests. But the barrier layer composition is believed to be ultimately quite complex due to ring opening crosslinking and other side-reactions. In a few instances. Of these. polyalcohols. water fluxes of up to 23 gfd (99. This membrane was discovered by Cadotte. followed by an oven curing step at 125’to 140°C.5% synthetic seawater at 25°C was 20 gfd and 99.500 psi test on 3.6% salt rejection) at 800 psi on simulated seawater” and 50 gfd (98% salt rejection) on 0. the 20. isopropanol was eventually added as a cosolvent to prolong the usable life of the dip-coating solution. Various water-soluble additives were also evaluated for their effects on formation of the sulfonated polyfurane resin and its subsequent reverse osmosis performance.000 molecular weight polyethylene oxide was singularly beneficial in increasing membrane flux quite significantly without disturbing salt rejection. The most outstanding property of the NS-200 membrane was its extremely high salt rejection.5% sodium chloride at 250 psin were achieved. generated a black sulfonated polyfurane resin that was embedded in the surface of the polysulfone. salt rejection levels of 99. The chemistry of the NS-200 barrier layer in its early stages of formation is believed to be approximately as shown below. polycarboxylic acids and polyethylene glycols. and partially decolorizing in exposure to the bleaching effects of sunlight. When the NS-200 membrane was initially discovered.9% were rountinely observed.69 and was characterized in government-sponsored research proThe NS-200 membrane in its optimized formulation71 was pregrams. The freshly prepared membrane is glossy black in appearance. The high oven temperature.
Membrane instability was believed to result from hydrolytic cleavage of sulfate and sulfonate esters in the sulfonated polyfurane resin. in addition to drying the membrane. and may be mostly absent in the final membrane after washout.

000 psi. but at extremely low fluxes (0. has many similarities to the NS-200 membrane. the extremely high salt rejections characteristic of NS-200 were obtained.Thin Film Composite
Reverse Osmosis Membranes
335
mechanism that was not fully considered during attempted development of the NS-200 membrane. When furfuryl alcohol was added as a comonomer to the THEIC. In addition.9% rejection and 12 gfd flux under seawater test conditions at 1.062. ” The cross section of this membrane. PEC-1000 membrane). The composition of the PEC-1000 membrane has not been specifically published. A typical patent example of membrane fabrication uses a water solution of 1:2:2: 1 weight percent THEIC:furfury1 alcohol:sulfuric acid:dodecyl sodium sulfate. The chemical structure of the PEC-1000 membrane is probably quite complex. This membrane. may involve some of the following types of chemical structures.366.76 This patent describes membranes formed by acid-catalyzed condensation of the monomer 1. showed 99. The barrier layer is 300 angstroms thick.25% NaCI.0 gfd).3. 570 psi.+ 0 a2(etc)
JY4~ ++CH2CH2 . it is covered by U. This membrane.d0 Ct. Further studies by other groups attempting to commercialize the membrane were unsuccessful in solving the NS-200 membrane instability prob72-74 lem. 2
OH
H2Xb
hH2CH20H
-I _CH2C”2\N a p+C”2 -n. under high magnification in a scanning electron microscope is quite similar in structure to the NS-200 membrane. for example:
HOCH2CHz‘N A &. This membrane has been described by Kurihara and coworkers as a thin film composite membrane in which the barrier layer was formed by an acid-catalyzed condensation reaction on the surface of microporous polysulfone.~O
ti~CH*CH20H +
UC. 25”C). However. Patent 4. developed by Toray Industries.5-tris(hydroxyethyl)isocyanuric acid (THEIC).CH2CH2Y
To
H2CH20-
. while the high organic rejections characteristic of the isocyanurate moiety were retained.5 to 1. deposited on microporous polysulfone and cured at 150°C for 15 minutes.S. possessing a thin active layer 100 to 300 angstroms thick. but which may have been an important factor (cf. By itself or in combination with additives such as sorbitol or polyethylene glycol. in its early stages of formation.
PEG1000
MEMBRANE
The PEC-1000 membrane. Organic rejections were very good. water fluxes were increased tenfold. this monomer leads to membranes with 95 to 97% sodium chloride rejection (0.

The type of display in Figure 5.5%.9
l
Hollow fiber
membranes
Thin
film
camposice
______________ Partial PT-30 membrane double-
membranes
stage process
98 97
_ membranes
.2 Permeate
. 2s”c
h
99. and maintenance of high salt rejections over a pH range of 1 to 13 (in short term tests).shows a comparison of several types of commercial reverse osmosis membranes in terms of salt rejection and permeate flow rate under seawater test conditions (35. If
r
f
3. This chart emphasizes the capability of PEC-1000 to provide complete singlestage seawater desalting. PEC-1000 spiral elements operated at 35% recovery produced a permeate having an average salinity of only 220 ppm. indeed. well below WHO standards.16. seavater
800 99. has been found to suffer damage from dissolved oxygen in the feedwater. This has been alleviated by using sodium bisulfite as a feedwater additive. Figure 5. but for organic chemical concentration and seawater desalination applications. the high rejection characteristics offset this disadvantage.000 ppm seawater (equivalent to Red Sea salinity). 79rs0. Average salt rejection was 99. In a test at Toray’s Ehime desalination test facility on 42. BOO psi. such as Kamiyama and coworkers8’ It should be noted that flux is shown on a logarithmic scale in Figure 5. and this presents a visual bias to the reader.5 Flow Rate
1 (m’/m2*
2 day)
3
5
Figure 5. 25OC).16 has been adapted and altered by others.3
. high organic rejections.
.16.57.16: Comparison of desalination performance of different commercially available membranes for seawater.77 The membrane is not chlorine-resistant.7a Flux of the PEC-1000 membrane is low relative to polyamide composite membranes.05
.5
pst. adapted from Kurihara. and.000 ppm.336
Handbook of Industrial Membrane Technology
Basic characteristics of the PEC-1000 membrane include very high salt rejections.l
. and by use of a vacuum deaerator as a pretreatment step in a large seawater plant design.

Thin Film Composite Reverse Osmosis Membranes

337

the flux axis were changed to an arithmetic scale, the flux advantages of polyamide composite membranes would become visually more evident. In the case of seawater desalination plants, two stage systems become economically feasible at the higher fluxes. In tapwater and brackish water purification systems, absolute flux is normally more important than small differences in rejection. In terms of organic rejections, PEC-1000 membrane shows the highest values among all commercial reverse osmosis membranes. Table 5.6 lists rejection data for a variety of organic compounds. In most cases, these were measured at solute concentrations of 4 to 5%. which represents a severe test protocol. Organic solute rejections determined for other commercial membranes were typi-

tally at solute concentration levels of 0.1 to 1.0%. Figure 5.17 gives a capsule illustration of the differences in organic rejections for four commercial membranes: PEC-1000, FT-30, Permasep 8-9 polyamide and asymmetric cellulose acetate. The superior organic rejection characteristics of PEC-1000 are quite evident in this graph.

Solute
Rejection (“/.I

60 50

PEC-1000 FT-30

composite

composite aramid
cellulose

Asymmetric
Asymmetric

(hollow
acetate

fiber)

Figure 5.17:

Comparison

of organic solute rejections of various membranes.

SULFONATED

POLYSULFONE

MEMBRANES

Sulfonation of aromatic polymers has been explored as a method to produce hydrophilic polymers with water permeability and salt rejection characteristics. These have been of interest because of their potentially high degree of chlorine resistance. The use of sulfonated aromatic polymers for reverse osmosis membranes began in the late 1960’s with the work of Plummer, Kimura and LaConti of General Electric Company. 82 Polyphenylene oxide [poly(2,6-di-

Thin Film Composite Reverse Osmosis Membranes

339

methylphenyleneoxy)] were cast from a chloroform/methanol solvent. The degree of sulfonation was limited so that the product would be water-insoluble. Later, composite membranes of this polymer were prepared by meniscus coating onto Celgard microporous polypropylene film and eventually onto microporous polysulfone sheeting.s3 In the latter case, a solvent blend was developed for the sulfonated polyphenylene oxide, consisting of nitromethane/methanoI/butanol, which did not dissolve the polysulfone substrate. While this membrane appeared useful for several difficult applications involving highly acidic or high temperature feedwaters, its stability and performance characteristics were not sufficient to result in its commercialization. After the initial work on asymmetric sulfonated polyphenylene oxide membranes, sulfonated polysulfone was subsequently examined by several research groups. However, the appearance of patents in 1973 and 1977 by RhonePoulenc~r8’. covering asymmetric membranes of sulfonated polysulfones appeared to dampen such research efforts. Also, a basic problem with sulfonated polysulfone was the fact that salt rejections fell rapidly when a critical level of sulfonation was exceeded, but water fluxes were too low until this sulfonation level was exceeded. Beginning in late 1976, Cadotte and coworkers developed a composite sulfonated polysulfone membrane, using a slightly different approach.46 Polysulfone was sulfonated to a degree sufficient to produce complete water solubility. Solutions of this polymer dissolved in aqueous or water-alcohol media were coated onto microporous polysulfone sheets, then oven-dried at 100 to 14O’C. This temperature treatment served to insolubilize the sulfonated polymer. Intramolecular sulfone crosslinks were assumed to be formed, though this was never specifically proven. Additives such as polyols and polyphenols could be added for additional crosslinking. These and other sulfonated membranes often showed good rejection toward dilute salt solutions such as tapwater or 0.1% aqueous sodium chloride. At increasing salinities, however, salt rejection levels showed a rapid fall off. The composite sulfonated polysulfone membrane described by Cadotte, for instance, exhibited 80% or less salt rejection in seawater reverse osmosis tests. In sulfonated aromatic polymer membranes, therefore, Donnan ion exclusion appears to be an important contributor to performance. In contact with dilute feedwaters, these strongly anionic membranes exclude feedwater anions from entering the membranes by charge repulsion effects, thus inhibiting salt passage. In more concentrated salt solutions, the charge repulsion effects are shielded by the high ionic activity of the feedwater. A sulfonated polysulfone membrane with commercial potential has been developed in the form of a hollow fiber by workers at Albany International Corporation.86r87 In a 5,000-hour test on 3,500 ppm brackish water at 400 psi, this membrane exhibited 98% salt rejection at 1 gfd. Flux and salt rejection remained constant even with addition of 100 ppm chlorine. In a 12,000-hour test on seawater at the Wrightsville Beach Test Facility, this membrane exhibited 98+% salt rejection and an average flux of 1.5 gfd at 1,000 psi operating pressure. Thus, it is possible to make a high rejection membrane from sulfonated polysulfone. Flux was rather low in this particular case, but suitable for hollow fiber membrane use.

Plasma polymerization has been explored in considerable detail as a potential route to composite reverse osmosis membranes. Yasuda has written an excellent review covering the fundamentals and progress of plasma polymerization through 1974.s8 Advances since then have not been particularly promising from a commercial standpoint, with the exception of the Solrox membrane (see below). Monomer polymerization by gas plasma techniques is quite complex, involving considerable fragmentation of vaporized monomers, and considerable crosslinking and grafting of deposited films. Vinyl unsaturation is not necessarily required for monomers to be polymerizable. Plasma polymers generally have short chain lengths, coupled with a high degree of branching and crosslinking, and containing a high concentration of residual free radicals. Monomers containing oxygen-based functional groups (such as carboxylic acids, ketones and esters) lose much of their oxygen content during polymer deposition. Nitrogen-containing monomers such as amines, on the other hand, show high levels of nitrogen incorporation into the deposited polymer films. Nitrogen gas and water vapor, added to the gas plasma of an organic monomer, tend to show finite rates of incorporation into the polymer deposits. To obtain good reverse osmosis characteristics, sufficient polymer must be deposited to “plug” the holes on the surface of a microporous substrate. Substrates having surface pore sizes of 200 to 300 angstroms appeared to be optimum in this regard. Early work by Buck and Davar” examined the plasma polymerization of several monomer systems. Best results in terms of reverse osmosis performance were achieved with vinylene carbonate/acrylonitrile and vinyl acetate/acrylonitrile. Wydeven and his coworkers examined the plasma polymerization of several amine monomers and obtained the best reverse osmosis membranes using allylamine.g0r91‘Reverse osmosis performances of 98 to 99% salt rejection and 4 to 8 gfd flux were achieved at test conditions of 1.0% sodium chloride, 600 psi, 20°C. ESCA spectra showed the nitrogen groups in the plasma-formed polymer to be nitrile or imine groups, but not amine groups. Oxygen was also present in the ESCA analysis. Elemental analysis showed a membrane stoichiometry of CsHs.sNo.&o.~. Yasuda obtained membranes with good reverse osmosis properties by polymerizing gas plasma mixtures of acetylene/water/nitrogen and acetylene/water/carbon monoxide.“r9j Elemental analysis of the acetylene/water/nitrogen membrane was not given, but it was likely approximately equivalent to the allylamine-derived composition. Yasuda has observed membrane performance levels as high as 99% salt rejection and 38 gfd, tested on 3.5% sodium chloride at 1,500 psi. To date, no commercial examples of these membrane compositions have appeared in the marketplace.

SOLROX 0200 MEMBRANE Sano and coworkers at Sumitomo Chemical Company developed the Solrox

Thin Film Composite Reverse Osmosis Membranes

341

membrane, which is technically a cross between asymmetric and composite An asymmetric polymer membrane is formed membrane technologies.B4* from a polyacrylonitrile copolymer. This membrane contains discrete micropores, being more like an ultrafiltration membrane rather than a reverse osmosis membrane. It is dried, then exposed to a helium or hydrogen gas plasma. The plasma treatment forms a tight, crosslinked surface skin on the asymmetric polyacrylonitrile support. Some vaporization and redeposition of the film components most likely occurs in this step. When placed on test, this plasma-modified membrane develops good salt rejection and flux in about 24 hours of continuous use. Analogs were also prepared using an organic monomer such as 4-vinylpyridine in the gas plasma. Data in a 1979 patent show a water flux of 87 gfd at 140 psi for the unmodified asymmetric film substrate.” After plasma treatment, this film exhibited 11 gfd and 98.3% salt rejection on a 0.55% sodium chloride feedwater at 700 psi. Table 5.7, taken from a July 1982 technical bulletin by Sumitomo Chemical, lists performance data for a series of Solrox membrane grades that can be made by this process. Of these, the Solrox SC-0200 has been manufactured in tubular form on a semicommercial basis and used within the Sumitomo organization on organic chemical concentration applications. Table 5.8 lists organic rejections for the SC-0200 membrane. These organic rejections, while greatly exceeding the capabilities of cellulose acetate membranes, are not appreciably different than rejection levels exhibited by aromatic polyamide membranes (FT-30, Permasep B-9). Nor do they match the organic rejection characteristics of the PEC-1000 membrane. The Solrox membrane is not resistant to chlorine, and its water flux is somewhat low (about 25 gfd at 650 psi net driving pressure). Consequently, it has not become a significant competitive membrane in the world marketplace.

The emphasis in this chapter has been on commercially significant membranes. Historical technical events leading to the successful development of these membranes, and information on closely related membranes appearing in the technical literature have also been presented. Other composite membranes have appeared in the technical and patent literature, that have been omitted for lack of time and space. Perusal of the U.S. government-supported research project reports published by Cadotte, for instance, will reveal composite membranes based on polyacrolein, styrene-maleic anhydride condensates, metaphenylenediamine-formaldehyde polymers, hydrophilic aromatic polyesters, sulfonated polymers, terephthaloyl oxalic-bis-amidrazone polymers, and crosslinked polyvinyl alcohol derivatives. A few other interesting examples are given below, illustrating the versatility of the composite membrane approach. Strathmann prepared an all-polyimide composite membrane-both bottom and top layers.” A microporous asymmetric film of the polyamic acid intermediate was cast by quenching in acetone, then dried and thermally cyclized to the polyimide at 300°C. The microporous polyimide sheet was then overcoated with a dilute solution of the same polymer, which was allowed to evaporate to give a 300-angstrom-thick coating. This was also cyclized to the polyimide to generate a fully solvent resistant reverse osmosis membrane. Taketani and coworkers at Teijin have sought to produce a stable flux polybenzimidazolone (PBIL) membrane by producing it as an ultrathin coating on

Thin Film Composite

Reverse Osmosis Membranes

343

The key to this effort was the discovery that the microporous polysulfone.g* PBIL polymer was soluble in water containing 5% or more of a water-miscible aliphaticamine. Thus, thesolvent sensitivity of polysulfone could be sidestepped. Koyama, Okada and Nishimura prepared composite reverse osmosis membranes containing an interpenetrating polymer network of polyvinyl alcohol with poly(styrenesulfonic acid).” The barrier layer was insolubilized by heating at 12O’C for 2 hours. Salt rejections in the low 90’s and rather low fluxes were observed. The long time length of heat-curing at 12OOCwas probably a factor in producing low fluxes. Kawahara and Yasuda obtained composite reverse osmosis membranes by acid-catalyzed crosslinking of methylolated poly(4-vinylphenol) coatings on polysulfone substrates.rea

ADVANTAGES

OF THE COMPOSITE

MEMBRANE

APPROACH

The numerous examples of composite membranes that have appeared in the technical literature exemplify the versatility of this approach. To the chemist attempting to develop new reverse osmosis membranes, the composite approach has three major advantages. First, large quantities of monomers or polymers are not required to fabricate the barrier layer. At surface thicknesses of 200 to 2000 angstroms, roughly 0.02 to 0.2 grams of barrier layer polymer per square meter of membrane is required (excluding excess chemicals used in membrane fabrication). These are orders of magnitude smaller than required in asymmetric membrane manufacture. Thus, very expensive chemicals are not a deterrent to composite membrane fabrication. Furthermore, one avoids the need for expertise in producing large quantities of high polymers in a reproducible manner. Small companies that are not basic in plastics production can emerge and compete effectively in this membrane market. In this context, only two polymers have ever been used on a large scale in asymmetric membranes: cellulose acetate and Permasep B-g/B-IO aramids. The former polymer predates the era of reverse osmosis membranes. The latter polymer has been used in hollow fiber membranes for 15 years. Attempts to bring other new polymers into asymmetric membrane production have been few (PBIL, PBI, polypiperazineamides), generally without particular success. Second, insoluble crosslinked barrier layer compositions are possible, and, in fact, are almost universal in the composite membrane approach. Optjmum reverse osmosis performance and chemical stability can be achieved, in part, due to preparation of crosslinked compositions. This is readily possible by the composite membrane approach, but not so simple by the asymmetric membrane appreach. The PA-300, FT-30, and PEC-1000 barrier layer compositions, for example, are simply not feasible to prepare by asymmetric film casting techniques. The composite approach, therefore, is far more versatile. Water flux through reverse osmosis membranes is considerably dependent on the hydrophilic character of the barrier layer. In the composite membrane approach, highly hydrophilic barrier layer compositions can be used, suitably insolubilized by crosslinking. To a large extent, water flux and salt rejection can be controlled by the type and extent of crosslinking. Third, each layer of a composite membrane can be optimized. The barrier

344

Handbook of Industrial Membrane Technology

layer can be developed to display single pass seawater desalination characteristics, on the one hand, or to selectively pass monovalent salts, on theother hand. The microporous substructure can be optimized for mechanical strength. The combination can be optimized for maximum water permeation rates. These advantages are reflected in the intense activity in filing patents on composite reverse osmosis membranes, and in the growing selection of commercially available types. Because of these advantages, many new examples of such membranes are likely to reach commercial status in the future, whereas very few new asymmetric membranes can be anticipated.

Formation of a concentration boundary layer in front of the membrane surface or within the porous support structure reduces the permeation rate and. the product quality as well. As laminar flow is present near the membrane surface. For this reason. At steady state conditions. This has been tested experimentally for a number of cases. Figure 6.. to be independent of the permeate flux V. specifically on flux. in most cases. at least with a good approximation. Consideration of only the transport mechanisms in membranes.e. the concentration profile is calculated from a mass balance as ‘I. the retained components must be transported back into the bulk of the liquid. At steady state conditions.1 shows how a concentration boundary layer (concentration polarization) forms as a result of membrane selectivity./ (1) ‘2-‘3 c1-c3 where k is the mass transfer coefficient which can be assumed. i. is based on a concentration gradient.6 Process Design and Optimization
Robert Rautenbach
INTRODUCTION-MASS The Local Mass Transport
TRANSPORT
AT THE MEMBRANE
SURFACE
It is understood that the economical success of any membrane process depends primarily on the quality of the membrane. the analogy between heat and mass transfer is valid and k can be calculated using the well-known heat transfer equations. For reverse osmosis.. this backflow is of diffusive nature. selectivity and service lifetime. however. lead to an overestimation of the specific permeation rates in membrane processes. will in general.’ = e k
9
349
.

leads to Pe numbers of less than unity. Figure 6. for this reason. Here. in combination with the permeation rates reported in literature. the error made in calculating the solvent flux V. this is impossible if a concentration profile exists within the porous support structure. Furthermore.02JRe7/8!k”L
400 . these diagrams explain why. in the case of gas permeation. turbulent flow is more advantageous. and the salt flux J. the resistance due to a boundary layer can be neglected: the diffusion coefficients of gases. cannot be influenced significantly by the concentration of the liquid on the product side.3. in contrast to separation processes in the liquid state. = c. in liquid systems. the local concentration affects the (local) permeate fluxes of the mixture components and.3:
number
Re
Concentration polarization in turbulent flow. Influence of the Asymmetric Structure of Membranes While it is possible to enhance mass transfer in the boundary layer by improving the flow conditions. Again. the concentration c3.\
Reynolds
Figure 6.2 and 6. For reverse osmosis.
According to Figures 6. According to Figure 6.4 shows the predicted concentration profile in the porous structure of an asymmetric membrane. is only 2% or 5% respectively.d
! -
!! Sh= ~.).Process Design and Optimization
351
OS “I
-
1
Pex-fj-
VI7I. which affects permeate flux and product quality. if the effect of cs is neglected. Even in the extreme case of local concentration on the permeate side being equal to that of the liquid on the high pressure side of the membrane (c. the type of flow for gas permeation on the permeate side is
. the situation is different in the case of gas permeation.4. Conditions are nearly always favorable for the permeate to flow unhindered. for example 02/Nz are lo4 times larger than those of dissolved components in liquids such as NaCI/l$O which.

5 shows theoretical results of Blaisdell and Kammermeyer.352
Handbook of industrial Membrane Technology
much more important than for liquid separation processes such as reverse osmosis.2A .
Concentration
profile in porous substructure of asymmetric mem-
0. feed
~membrone-----_tcpermeate
porous subslucture
membrane
thickness
Figure 6. Figure 6.
t
1
1
I
I
0.4: branes.complete
mixing
stage-cut
0
Figure 6.2 Mol fraction of oxygen is plotted against the yield during fractionation of air for different flow configurations and cases of mixing.29.5:
Effect of flow configuration in gas permeation. The diagram demonstrates that (1) countercurrent flow provides the best results and (2) marked differences exist between the perfect mixing and plug flow.
.

.= &[p”+~pT. C. . the membrane is of tubular design with feed-flow inside.+ P %:..(j VW*.
3
I
for mass transfer
Sh = y
= ShlRe.APV
.I . = K...
-P.
(2)
(3)
For these reasons.. = A... d
l-~.( =TX..
salt concentration balance mass for a length
C) = Ir VW element v.V. no influence of the porous support-layer) are:
mass transfer concentration equation at the membrane polarisation e surface I evw’.-c.‘&
.“.. this has to be taken into account for one or both sides..n.n. the overall permeation rate and the product quality of a membrane channel have to be calculated by integrating over the length elements employing.
Sc)
mass transfer permeate salt flux flux
in the membrane V.
balance balance (salt)
material
‘jr.~(1-5
x. (1) the mass balance (2) the material balance (3) the energy balance.)
.
CI. Decrease of the mean velocity ( + increase of concentration polarization + decrease of permeate flux and increase of salt flux). If.. the following has to be expected along the tube-axis: (1) Decrease of the transmembrane pressure-difference as a consequence of the drop in pressure due to friction ( + decrease of permeate flux).Process Design and Optimization Change of Conditions Along the Membrane
353
As a consequence of the permeate flux.IAp-blcz-c. AP ‘P. Increase of the mean concentration of the feed solution ( + increase of osmotic pressure + decrease of permeate flux and increase of salt flux).~.LAX . in addition to the equations discussed above.*.)) 3s =
WC. for example..
The equations necessary for the calculation of tubular RO-membrane channels (flow of permeate unhindered..9
energy
balance(p=constl
pm+.Cl. the pressure gradient as well as the mean velocity and concentration will vary along the membrane surface-in principle at the feed side as well as at the product side.Pn
LAX. Depending on module design..

the industrial configuration of membranes. Optimization procedure and some of its results will be discussed for only one module configuration. The most important designs are the plate and frame..
MODULE
CONCEPTS AND DESIGN
The module. (3) price of module and (4) the possibility of cleaning the membrane. at the same time. the hollow fiber module-and for only one application-RD.e. a number of different modules have been designed. Points of major importance are: (1) flow conditions along the membranes. The Hollow Fiber Module Hollow fiber modules contain very fine fibers forming asymmetric or symmetric membranes and capable of withstanding pressure differences (high-pres-
. Figure 6. the increase of local and mean product concentration cs and &.364
Handbook of industrial Membrane Technology
This set of equations has to be solved numerically. It clearly demonstrates the decrease of permeate (mass) flux J. (2) ratio of membrane area to pressurized vessel volume. the spiral wound and the hollow fiber module. Depending on the process application. has to meet a variety of requirements which are sometimes even contradictory.6:
Calculation of a tubular module. one or the other of these requirements is of primary importance and.6 discusses the results of such a numerical calculation of a long tubular RO-membrane channel for seawater desalination.
0
‘0
20
LO 60 Number of tubes
80
100
Figure 6. i. and. for this reason.

the membrane area per unit volume decreases.. However.8 clearly indicates that hollow fiber modules for reverse osmosis should be made as short as possible.
According to Figure 6. Equation 8 is discussed in Figure 6. For some simplifying assumptions. the specific yield of the fiber bundle is given by:
$2$&d&. Neglecting (1) the radial pressure losses in the feed and (2) the influence of osmotic pressure on flux. at the same time.10:
Optimum design of hollow fibers.10. It should be noticed that a simple correlation between yield of the single fiber and yield of the bundle can be derived in any case only if the concentration and the pressure in the shell (outside of the bundle) are considered to be constant! In principle.10. it can be easily discussed how fiber diameter and fiber length have to be chosen in order to maximize the “yield” per unit volume of the bundle.e. larger fiber diameters should increase the flux. an optimal fiber length exists for a given fiber diameter. the performance of a single fiber. v
(8)
4(1-E) d
3 3
tanh (HLI l+Hsls*tanh (HL)
For a fixed ratio of 2
I
= 2.
I
0
I
20
I
LO
fibre
I 60
diameter
I 80
di
1
plm
100
Figure 6. the concentration of the feed increases between entrance and outlet while the pressure of the feed
. i. Figure 6.358
Handbook of Industrial Membrane Technology
In theory.

11).Process Design and Optimization
359
decreases.
In the case of reverse osmosis. K. has to be varied according to the fiber arrangement. A numerical solution taking this into account can be achieved if the bundle is considered as a continuum containing sinks (Figure 6. As already indicated. While countercurrent flow improves the separation efficiency of hollow fiber modules in reverse osmosis only slightly. concentrate
fiber
bundle
on
permeate
Q39C30
1
1
feed solution Qla. = 30 is valid for fibers in parallel. (t-11 =
energy balance: (11) dp. as far
. In this case. (r) -a-i=-= K0 nFtv]
U-EIZ
Ej
.
1
(r)
K.
ClorlPISX
Figure 6. the following equations have to be solved: mass balance (pL = const): (9) material balance: = -& (r V. the situation is quite different for gas permeation. the relative flow configuration does not affect the performance to any large extent.11:
Balance element in the fiber bundle.

CASCADES There are cases where it is impossible to fractionate a given mixture in one
. product 1 (permeate I
4
f i
I
enriching
section
I I I
feed
I I I I I
stripping section
product 2 (concentrate Figure 6.12).. membrane area per unit length). for example. a module is described based on this principle. This column consists of a rectification and a stripping section and permits binary mixture to be fractionated into its components.e. countercurrent flow produces the best results. For a given geometry of the membrane fractionating column (length of each section.360
Handbook of Industrial Membrane Technology
as gas permeation is concerned.
The modules employed in a membrane fractionating column must be different from those used in cascades: in hollow fiber modules.4 i.12:
1
Membrane column for gas separation. the fibers cannot be U-shaped and must be open at both ends. the membrane column (Figure 6. A consequent application of the countercurrent principle led to the concept of Hwang. the product quality is determined by the flow rate of the gas being recirculated. In Reference 5.

a gas centrifuge. for example. i. Separation units of a separation process could be. that in plate-and-frame modules as well as in modules of the spiral-wound-type usually every block or pressure-vessel contains more than one separation unit! Stage. Furthermore. for example. the unit itself will be the stage of the cascade. the trays are still incorporated in one vessel. Separation units are considered part of a stage as long as they are connected only at the feed and/or retentate side! The elements of a stage can be arranged according to Figure 6. however. gas centrifuges and-last not least-membrane modules. The element of a cascade will be named “stage” and the stage itself might contain several separation units connected in series and/or in parallel. part of one unit. Definitions Separation Unit. These stages can be realized in several ways.
. the separation process will consist of several stages. In other cases. The elements of a separation process.Process Design and Optimization
361
stage to the desired degree of purity. processes consisting of more than one stage are known in gas permeation and for the production of boiler feed water from seawater. in a “tapered” fashion or in a “squared-off” fashion. cascades of hydrocyclones. by employing the countercurrent flow principle as much as possible within one unit like in packed towers for absorption or rectification or by combining definite separation units in the form of a cascade. a membrane module. cascades will be necessary if organic mixtures are to be separated by membrane processes only. In membrane technology. a tray of distillation column or the evaporator of a multiple effect plant.
Figure 6.13:
“Tapered” module arrangement.. The best known examples of this case are fractionating towers employing trays. The best known example of a membrane cascade consisting of a high number of stages is the Oak-Ridge plant for enriching U-235 in the gaseous phase (employing porous membranes).13 (this is sensible with respect to mass transfer in the boundary layer because of the decreasing volume flow at the feed side along the “axis” of a stage). for example. It should be kept in mind.e. In fractionating towers.

Yield is defined as the ratio of a certain (key) component in the product and in the feed:
(12)
Selectivity.“N
(Retentatel
Cascade without reflux. Yield.
.362
Handbook of Industrial Membrane Technology P(Product1 t
.
Cascade. Selectivity in general is defined as the quotient of ratios of conveniently chosen concentration measures.
PA’PB 1 Prod 5 /PA PBI Feed
= ‘A”BI Prod ‘A’“B 1 Feed
or
(14)
s2
WA”JB 1 Prod = -i--7?A B1 Feed
. the selectivity is defined as:
(13)
s. for example.. the partial pressure pi and the molar fraction xi or a mass fraction Wi. Convenient measures for a concentration are. A combination of stages where the permeate of a stage will be the feed of the next stage. Accordingly.14:
RN ‘.
yp
Rnlxn .bed&
Figure 6.

the operating line represents the mass and material balances. Another entity which is often employed because of its simplicity is defined as:
(16). however. consequently. sometimes called “split-factor”. for the graphical determination of the number of stages of a cascade. This method has been developed for the design of distillation columns by MacCabe and Thiele and should be well known.
or xn =
x
=-'AI Permeate xAj Retentate
=
yA x A
YAPI module p ermeate.
(15)
.
1
=
Qi
(PI
xi
-
P2
Yi
) membranes.Process Design and Optimization For the permeation of ideal gases.e. Equation
describes the transport mechanism (in sorption-diffusion) 13 is reduced to QA s=-=. There will be. i. for all membrane processes. for example.Qiprxi is valid). stage n 'AR1 module retentate.
. ni .
n. “Cut-Rate” (Split-Factor). the equilibrium curve represents the thermodynamical vapor/liquid equilibrium.
(17)
or
(18)
Operating Line and “Equilibrium” Curve.. a difference between the cut-rate defined for molar flows and for mass flow. the equilibrium curve does not represent a thermodynamical equilibrium at all but will represent the separation characteristics of the module or that of the stage. the equilibrium curve can be calculated from Raoult’s law and the saturation-pressure curves of the pure components of the mixture. stage n
=-Yn xn
if characterizing the selectivity of the entire module. In all other cases.O<ol<l
QB
(in cases. The cut-rate. however. For all cases. where psvi << pr Xi and. In distillation. where
363
. is defined as the ratio of permeate flow and feed flow. Both terms are of importance for the graphical solution of a separation problem. For an ideal binary system.

there are four possible modes of operation: (I) Constant reflux R/P
(2) Constant cut rate and variable reflux (3) (4) Variable cut rate and variable reflux “Ideal” cascade operation (x. the operating line of a cascade follows from a mass/molar balance and a material balance between the top and a certain stage n of the cascade. Constant Cut Rate. = v = const. Constant Reflux. Figure 6.364
Handbook of Industrial Membrane Technology
Cascades Without Reflux Such cascades are only sensible in cases where the retentate is practically worthless.R = x.15. Variable Reflux.p). = 0
Rn -
is developed for several stages. Only a detailed calculation of the specific separating costs will lead to a sound decision on whether a separation process should operate with or without reflux. the number of stages necessary to achieve a certain product quality can be easily found according to Figure 6. has to be weighed against the higher capital costs of reflux cascades.
= g
Yn+l
= v+l
%
1 G-i
y1
The operating line is a straight line with the slope --$ If the “equilibrium curve” is known.14 illustrates the principle of a cascade without reflux. For constant reflux R = const resp.
= 0
P
n+l
yn+l -
Rnxn
P. however. _ t. only the first two cases will be discussed because with modern computers at hand.
y. for example.16 by a step-by-step graphical procedure. In principle. is operating an electrolysis-heavy-waterproduction-plant where the deuterium is separated in a cascade without reflux (at least in the lower part of the cascade). Norsk-Hydro. Reflux Cascades The inevitable losses of product with the retentate of cascades without reflux can be avoided by recycling the retentate according to Figure 6. (22) P n+l When the molar balance: P. This advantage. the molar balance: P and the material balance: (20) result in: (21)
V
n+l
-
Rn -
p.
Here. for these two cases only a graphical solution will be of value. In general. + 1. R/P. a pattern can be noticed if the (constant) cut rate is introduced:
.

i. countercurrent flow Partial (incomplete) mixing of feed and permeate (Breuer Case’ ).
(P. the equilibrium curve in this case is influenced by the ideal separation factor So and the pressure ratio r6.Process Design and Optimization The “Equilibrium Curve”
367
As mentioned above.
(p*(l-x. the “equilibrium curve” in the case of membrane cascades is not at all a thermodynamic equilibrium but will represent the separation characteristics of the module or of the stage..
x.‘-I-
QB/QA leads to:
-‘pn
= L
o
-. for example.lO~‘l Complete Mixing of Feed and Permeate.18). five different patterns are possible: (1) ~~z.or countercurrent flow. plug flow at the feed side and permeate flow orthogonal to the membrane without mixing (Weller Steiner Case 2 6 Naylor-Baker’)’ (3) (4) (5) Parallel plug flow. yn yJ6 ( Ynx n -1 -(l.
p2(1-
Combining of both equations and introducing 6 = ps/pr and o = (30) n .e.
Here. This separation characteristic is in general a function of: (1) the selectivity of the membrane (2) the fluid dynamics in the module (3) the driving force (4) the concentration level (because the membrane selectivity is or might be concentration dependent) (5) the flow pattern in the module. A detailed discussion of the other cases can be found elsewhere.r?$e mixing at both sides of the membrane (Weller Steiner
(2) Cross-plug flow. In this case (Figure 6. co. cocurrent flow (Blaisdell-Case’) Parallel plug flow. the equilibrium curve simplifies to: (31)
). For large pressure ratios.?2
> -
y*)
Yn))
= Q. .(l-
Y * Y n)
=
Q.n
.c^
As Equation 30 indicates. With respect to flow pattern. only the first case will be discussed. the material balances for a binary mixture are:
P*
(29) P.

05
---
I
1. At present. The preferably permeating component can be withdrawn at the top of the column.
---
a = 0.5
xn
Separation characteristics (complete mixing).
. However. is restricted to cases where concentration polarization is negligible. for this reason.) Membrane columns utilize the maximal driving forces because of the countercurrent flow at both sides of the membrane.2 a =0.0
0
Figure 6. all modules for membrane columns described in the literature employ hollow fiber membranes.368
Handbook of Industrial Membrane Technology
T
Figure 6. the slower permeating component of the mixture at the bottom.20) and. The preferably permeating component is depleted on the high-pressure-side in the direction of the flow. this concept must fail here-at least for asymmetric membranes because of the concentration polarization within the support layer.18:
pnl
Yn
Complete mixing.19:
0.12-14 The membrane column consequently utilizes the countercurrent flow principle (Figure 6.
Membrane Column The membrane column can be an alternative to membrane cascades. (It has been suggested that the membrane-column concept be applied to the separation of liquid mixtures.

Figure 6.Qi
-1
11
( \-_. The lower limit is set by the condition that diffusive axial backmixing must be negligible with respect to the convective mass transport.
(2)
The separation process is determined (among other parameters) by the reflux ratio.1 l-Oj/Qi
l-Xi. on the other hand.
lo. a discussion of Equations 40 and 41 is useful for a first estimate of the column length and the operating conditions.
Membrane
column. For these reasons. Xia l-Xi
-1
1
Xia
+1_xla
Qj/Qi l-Oj/Oi
1
Xi -Xia
)
-..(l-2)
&t I
l-Qj/Ql
Xi l-xi.
l-Xi
I-d.
i
=
ka
Equation 36 can be integrated:
l+Oj/Qi Qj /Oi
h
a
x. the volume-specific yield (yield per unit-volume of fiber bundle) should be high.
material
balance around a fiber end (total
With the boundary conditions: (39)
z = 0:
xi
= xi
a .22 indicates that a specific molar flow rate (molar flow per fiber) as low as possible should be chosen.
l-Xi. the calculation with total reflux leads to the minimal necessary column length. On the other hand.. For a given column length.. Similar to fractionating by distillation. Column length depends on the recycled molar flow rate f&. Equations 40 and 41 indicate the maximum possible purities of top and bottom product.Process Design and Optimization
371
ki
Figure 6.21: reflux).
(41)
ti = ti.
. -AT
l-Oj/Qi
(40)
z=nd Qj p.

are the reason why hollow fiber modules for the membrane column are relatively long (L/d. Figures 6. The concentration profile is very similar to the profile calculated by a numerical solution of the set of Equations 37 for the separation of C0JN2 (Figure 6.23 shows the concentration profile of a membrane column as calculated by Equations 40 and 41 (total reflux).24. clearly indicate that a membrane column is especially effective in the stripping section. taking into account the pressure losses and a finite reflux ratio. the optimization of a membrane column must consider molar flow rate &.
column. For the enriching section. which have been only shortly discussed here. The mutual dependencies.-20. at the most. absorption.24).
(recycled
molar flow rate)
In general. whether a membrane column or a reflux cascade is superior cannot be answered in general but must be evaluated for every individual case.000) compared to hole low fiber modules for the separation of liquid mixtures.23 and 6. The question. Figure 6. According to our opinion”.372
Handbook
of industrial
Membrane Technology
10
0. two-stage cascades seem to be competitive compared to proven processes like pressure swing adsorption.0
21) m L
length
of reflux
Figure 6.
. the necessary membrane area and the reflux ratio (pressure losses) are increasing rapidly. only one stage or. or
low-temperature distillation.22: Membrane on column length. reflux ratio and the pressure losses. It seems doubtful. especially for high concentrations of the preferably permeating component.1 Fiber (Column) influence
1. It should be emphasized that a membrane column represents a multi-stage process just as a reflux cascade. whether multistage membrane processes will be often employed. however.

The feed pretreatment is simple: sand-filtration (well water) followed by a settling tank. the primary energy consumption of seawater desalination can be kept Iowl which is one of the reasons that thermal processes will always have their place in seawater desalination although the relatively rigid coupling of power and water production is sometimes disadvantageous. This trend can be noticed at least partly in the following examples. RO plants are.1 summarizes the major op erating data of the plant. The modules which are successfully employed are of the “spiral wound” or the “hollow fiber” configuration. The energy consumption is approximately 95 KWh/to distillate in the form of saturated steam of a pressure of about 2. This very low registered figure includes the power consumption for the feed pretreatment and is a direct consequence of the installed energy recovery system (Guinard integrated turbo pumps). when compared to proven separation processes such as distillation. it is accepted technology to use well water from wells drilled at the seashore rather than surface water as in thermal desalination. Figure 6.3 KWh/to fresh water. Table 6.374
Handbook of Industrial Membrane Technology
PROCESSES Any application of membrane processes will be ultimately determined with respect to economic reasons. consisting of 10 trains of hollow fiber modules (BlO Permeators of DuPont).
Seawater At Desalination by RO
present. In general. Under favorable conditions. the economic service lifetime of the modules resp.25 shows the process flow sheet of the plant. For this reason.000 t/d. All distillation processes produce a distillate containing only 20 to 30 ppm salts (total dissolved solids). By this combination.000 t/d fresh water at a specific power consumption of 6.2 is a breakdown of cost. membrane processes will not replace conventional processes but. much more sensitive to insufficient feedwater pretreatment than thermal plants.2 bar and 4 to 5 KWh/to distillate in form of power for the pump drives. Modern distillation plants have unit capacities of about 25. Feedwater temperature is not a critical factor in this case (in general. all major plants are combined with power plants-the heat input section of the desalination units replacing the condenser in the Rankine cycle. produces 20. a combination of membrane processesand conventional processes will be the optimal solution. some data of the Malta facility shall be discussed. Table 6. in the main. the membranes is about 5 years and warranties for such a figure are given by manufacturers. acid dosing and cartridge filtration. however.” The entire plant. Strictly speaking. As an example of the present state of the art in RO-seawater desalination. the manufacturer limits the maximum temperature to about 30°C despite the fact that the membrane itself can withstand higher temperatures). Reverse osmosis (RO) will be the alternative to distillation. the following measures have to be considered for feed pretreatment of RO seawater plants:‘s
l
pH-adjustment by acid dosing against carbonate scaling chlorine dosing against marine life in the plant
a
. nearly all seawater desalination plants are distillation plants. Since low pressure steam can be utilized for heating.

000 1.26. i. In general. a flushing and cleaning system is necessary.11
0.e.e. The module arrangement in each stage.46
Chemicals
4.14 0.27).03 0.26) are not economical. The second stage can be equipped with a different type of membrane (high-flux) than the first stage and certainly must be operated at a different transmembrane pressure difference.000 5.26) for the improvement of the flow conditions in the modules and/or pumps between the module-banks of a single stage (DEP in Figure 6.01 0.250.376
Handbook of Industrial Membrane Technology Table 6. these requirements lead to a so-called “tapered design” (see Figure 6. a hybrid plant consisting of a reverse osmosis process as a first stage and a special crystallization process as a second stage seems to be feasible (Figure 6.04 2.2: Production Cost Summary Malta Seawater RO Plant
S/m’
S/Kgal 1.08
0.13) In any case. active charcoal filtration and/or sodium bisulfite dosing is mandatory if chlorine has been used in the pretreatment.1912o Total Desalination of Brackish Water When desalting brackish water far inland.20 0.62
0. much more than the distillate from thermal desalination plants. the designer faces on optimizing problem which can be solved numerically by a relatively simple computer program based on the general flow sheet indicated in Figure 6. the discharge of the concentrated brine may pose a serious problem. For the case where neither deep-well-injection nor a discharge into a river or canal is possible.
. especially in the proceedings of the IDEA conferences and of the symposia “Fresh water from the Sea”. Higher product purities-for chemical processesor for boilerscan only be achieved by a (2-stage) RO cascade.65 1.07
Furthermore..05 0. More details about seawater desalination by RO can be found in the literature.000. i.43
Electricity Permeators Labor Spare Filters Total Total costs Civil Total Operating Investment RO-plant works production L Overhead Parts
0.43
0. One result of numerical calculations which seems to be generally valid is that recycling pumps (ZP in Fig ure 6.01 0. For membranes which are sensitive to oxidizing agents. the number of modules in series and/or in parallel is determined for a given type of module by the desired yield of the stage and a necessity to maintain optimal flow conditions in the modules (concentration-polarization). Note that the product of any one-stage RO seawater plant contains about 300 to 500 ppm total dissolved solids..04 0.

at the inlet should be about 80.v&.26: General flow sheet of an m--stage RO cascade.27: crystallization.Process Design and Optimization SEAWATER
377
CONCENTRATE HDl
m. Re = h.
WATER
1 Cry s’ar’ion]
Solid Figure 6. STAGE
FRESH Figure 6.. for example ROGA 4160 HR. If modules of the spiral wound type are employed. Although the
. the Reynolds number. Total desalination of brackish
waste
water by a combination of RO and
Studies on three real brackish waters of different sources indicated that for the RO stage the optimum pressure will be between 40 and 50 bar.

30). For spiral wound modules (Figure 6.
. (2) Because of its high capital cost. Costs due to frictional losses and additional pumps are higher than the possible gain in permeate-flux by improved flow-conditions.e. the choice of the crystallization equipment is limited. of course. the agitated thin film evaporator..
to combine
IO 5 Number of modules in series
15
Figure 6. it is in general out of the WeStiOn agitated thin film evaporators to an energy SaVing multiple effect evaporation process. However. therefore. the brine of the RO stage. this type of equipment must lead to high specific treatment costs for two reasons: (1) An agitated thin-film evaporator is costly and limited iv: size.30). its influence on the minimum specific cost of permeate is nil in a wide range (Figure 6.29) as well as for hollow fiber modules (Figure 6.28: Influence of flow conditions in the modules on the specific cost of RO (tapered arrangement). the calculations clearly indicate that internal recirculations and/or repressurizing pumps between the module banks of a single stage will not pay off (see Figure 6. the lowest costs are obtained without recirculation (z = 0) and without repressurizing pumps (in Figures 6. The second stage of the hybrid-process is essentially a crystallization capable of crystallizing all solids dissolved in the feed resp. on the ratio of modules in series and in parallel.26).378
Handbook of Industrial Membrane Technology
Reynolds number has a significant influence on the plant structure.‘l Furthermore.29 and 6. Parts of these substances tend to heavy scaling and.28). A proven solution is. i.

but the fact that a rising brine concentration in the RO stage is equivalent to a decrease in feed flow for the crystallization step.31 comparing specific costs of different process combinations. For the RO stage. this is mainly due to the rising osmotic pressure. the specific treatment costs of each stage must increase with increasing concentration of the RO brine. results in smaller but specifically more costly crystallization units. it is easily understood that the optimum of the hybrid plant is determined by the maximum concentration which can be achieved with the RO modules! This result is demonstrated by Figure 6.Process Design and Optimization
381
One of the most important variables for the hybrid process is the concentration of the RO brine.31:
Tube Stage
Evaporator Elash
Water cost of alternative hybrid plants. In principle.
Capacity
ATF: Agitated Ihin
RO
Eilm
Evaporator
VTE: Vertical MSF: Multi
Figure 6. For the crystallization step not concentration per se.
. However. because the specific treatment cost of the crystallization is in any case higher by an order of magnitude than the specific treatment cost of the RO step.

combination
of RO and
Automotive Industry-Combination of UF.^^_a^
COllVENTIOt4AL PRETREATMENT Addatives vopour EVAPORATION
I
I
__---L___. The spent wash water of such a process contains. regardless of whether a VTE-seeding process. This is a direct consequence of the specific costs of the second stage. The process is designed to handle 10. (VTE = Vertical Tube Evaporator. etc.
Cleonlng
of flue llue gas dust
Feed
waler
treatment removal Calcium__.382
Handbook of Industrial Membrane Technology
In any case. the chances for such a hybrid process are very limited with respect to brackish water desalination.
. New Mexico. ATFE/VC = Agitated-Thin-Film-Evaporator in combination with Vapor Compression).
of power-plant
wastewater. MSF = Multi-Stage-Flash. that a process as described has been realized for the treatment of the effluents of a power plant (San Juan. Fractionating the wastewater into a concentrate containing the oil and the suspended particles and into a premeate containing practically no oil is possible by UF.
I
Permeate
Figure 6. the specific costs of fresh water for a 2-stage hybrid process will be 3 or even 4 times higher than for a single-stage RO process.32: Treatment crystallization. is usually a rinse. Scaling in the evaporator is prevented by the application of a (calcium-carbonate) seeding technique. The concentrate of the RO unit is fed to the crystallization stage.32). cutting oil and suspended matter and it is impossible to discharge it into the canalization without further treatment. axles. The mechanically pretreated effluents are fed to a RO unit. in addition to the detergents. Cross Flow Filtration and Evaporation for Recycling of Detergents and Process Water The final step in the production of automotive parts such as pistons. designed for a recovery rate of 80%. The situation seems to be different for wastewater treatment and it is not surprising.. designed as vertical tube evaporator combined with vapor compression. Because of these high treatment costs. here in combination with a seeding technique. a modified MSF-process or vapor compression is considered. USA).000 m3/d wastewater from the flue gas scrubber and the cooling towers (Figure 6.

tubular modules equipped with UF membranes of different pore-sizes and a microfiltration membrane have been tested.) As expected.33 because the permeate flux for steady state conditions is almost exactly the same for the three tested membranes (for comparison.8% suspended solids (metal.
50
30m/s. The savings in detergents are of prime importance and it must be the aim of process development to maximize the detergent concentration in the permeate without increasing the oil-concentration to an intolerable amount. its biological oxygen demand (BOD) is high. For two obvious reasons. the MF membrane produced the highest flux for detergents.33 have to be corrected for the same temperature and the same flow conditions). In this case.Process Design and Optimization
383
The detergent concentration of the permeate and. the oil concentration of the permeate is about 50 mg/S?. For ultrafiltration as well as microfiltration the permeate flux is controlled by a gel layer-not by the membrane itself.
@a
Membrane XT. the question arises whether such a permeate could be recycled: (1) a further treatment or further treatment costs would be obsolete and (2) the recycling would lead to substantial savings in process water and in detergents. etc. leading to further treatment costs if discharged into the canalization.). contrary to the UF membranes. dirt. In on-line experiments with a plant effluent containing 2% detergents. consequently. however. Because of th relatively low retention capacity for oil. where the oil concentration of the permeate is independent of the oil concentration of the concentrate up to the phase inversion concentration of about 41%.
33 bnr
0
10
20
30
60 min
Figure 6. (The experiments were carried out in the gear and axle production of Daimler Benz AG.
. the use of MF membranes is limited to a maximum oil concentration of about 10% in the concentrate. the data of Figure 6. 1% oil and 0.33: Gel-layer controlled flux in UF and MF (MF-membranes with and without backflushing). This must be concluded from Figure 6.

Ap=3.5%.Om2.S. of reclaiming metals which would otherwise be lost would be most favorable.
for the
The recovery of detergents of such a process has been determined to an average of 85%.34..3bor T=36*C. The optimal process is a combination of microfiltration and ultrafiltration according to Figure 6. Microfiltration is advantageous compared to UF because microfiltration membranes are available in the form of tubes withstanding outer or inner pressure of several bars which can be easily cleaned by reversing the flux for a short time. more than 200 RO plants are installed in the gal-
. In every case. Almost certainly evaporation will not be economical for the small capacities indicated in Figure 6. that the permeabilities of the individual components of a detergent differ and that the upgrading of the recycled permeate must be based on the selectivity of the process. an evaporation step must be included in the process. microfiltration alone will not be a solution for the described effluent-treatment problem since the retention capacity for oil at higher feed concentrations is too low. the gel layer can be removed by cleaning. the formation of a gel layer is reversible. At this concentration.34: Flow sheet for a combination of micro. the overall water recovery of the process is above 97.6m2. The limiting oil-concentration for the first stage. Y =S.6 m3/d 1 % Oil F=2. at the same time. In the U.Ap=2bor T=38OC. Since the oil retention capacity of the ultrafiltration stage is independent of feed concentration. It should be noted. Despite the advantages with respect to flux and yield of detergents. the final concentration of the retentate is solely determined by the phase inversion “oil/water -+ water/oil”.F=1. Y =3m/s
TO OR
COMBUSTION REFINERY
0. periodic cleaning and a higher average flux have to be weighed against the time for cleaning. Galvanic Industry-Treatment of Effluents
In the galvanic industry. is determined by the tolerable concentration of oil in the recycled product. However.and ultrafiltration treatment of oily wastewater. the retentate of the process still contains too much water if incineration or refining is considered. Therefore.16m3/d 10% Oil ULTRAFILTRATION .34 and a central evaporation station for the tentates from several production lines should be considered.Sm/s 1. A process which is not only capable of water treatment but. a treatment of the effluents is mandatory.24
MICROFILTRATION FEED F 11. RO is such a process.384
Handbook of Industrial Membrane Technology
Quite often. microfiltration.12 t-n31 d 90 %
Oil
1 EVAPORATION
1-1
Figure 6.

3 shows the (average) composition of the galvanizing bath.value temperature
.35:
Flow sheet of nickel plating. the permeate must be added to the first or second stage of the rinsing cascade.3allm8fl wa!er
37mgll
.04llmIn Delonlrlng 0.25
Evoporat~on Ltquid 0. Depending on the permeate concentration which.Process Design and Optimization
385
vanic industry and.3 60 =C g/l g/l g/l analyzed
brightening pH.04llmm losses
l0. Figure 6. these plants are very profitable because of the reclaimed metals. the evaporation losses of the plant are balanced by the deionized water necessary in the last stage of the rinsing cascade.6aIlmln 575mg11
5.6alImm
zadoomg /I
I
Figure 6.36 shows flux and selectivity (retention rate) of the modules.oLl/mln I_ I balh losses 0. inclbding RO for wastewater treatment. the recycling is complete.361Imm 14300mgll ’ Reverse 05nlos~* 5.
By recycling of the permeate into the first stage of the rinsing cascade and recycling the concentrate into the galvanic bath.3at/mm
Orogglng I Golvoniz!ng 2aoooomgII 9. Figtire 6. in turn. according to reported data. The modules are of the spiral wound type and the membrane material is cellulose acetate in this case.35 is a flow sheet of a galvanic production line completed by RO.Sml I 1027mgll c 1 407mgll 1 * O.. Table 6. Table 6. depends on membrane retention and on the required recovery rate of the RO unit.1?04 Ilmin
LOBrngll
1703 ltmin 9663mgA Reverse 0srnos15
11.3: Composition of Watts-nickel bath concentration total nickel nickel boric nickel sulfate chloride acid agent ratio 6 H20 6 HZ0 82 g/l 255 105 45 not 4.

36: selectivity. 2. Payback time of the RO units is given as 8. module flux and
Figure 6.4 years for a two shift operation and 1.386
Handbook of Industrial Membrane Technology
80
60
50
0
200
LOO
600
800 Time
1000
1200
1400
h
I -ii 600
200
LOO
600
800 Time
1000
12
Figure 6.
Treatment
of nickel-plating
effluents by RO.1 years for single shift operation.36 indicates the high retention capacity of the membranes for nickel and solubles whereas the retention for boric acid is not quite satisfactory (no explanation could be found for the concentration fluctuations in time).4 years for a three shift operation. An analysis of the RO units installed in the galvanic industry demonstrates
.

the evaporation losses of the bath are such that no additional evaporators for further concentration of the RO concentrate are necessary. Here. gas separation started with the recovery of Hz from the bleed of (high pressure) synthesis loops. Behind the scrubber. the permeate leaving at a pressure of about 70 bars.. At this pressure. employing in most cases a composite membrane “silicon/polysulfone” in the form of hollow fibers. the composite membrane differs from RO composite membranes: in gas permeation.Process Design and Optimization
387
that presently RO units are mainly installed in the nickel plating industry for two reasons: (1) the pH value of Watts nickel baths is between 3 and 5 permitting the installation of cellulose acetate membranes (cellulose acetate. the bleed is fed to a conventional separation unit first-a scrubber for recovering ammonia. gas permeation on a technical scale employs membranes of the sorption diffusion type. the NS 100 of Abcor Inc. Normally.. In all technical processes. At this temperature. Consequently. the material of the coating chosen (silicone) has a high permeability but a low selectivity while the membrane material (polysulfone) has a high selectivity (and a much lower permeability). Tokyo) which proved to be resistant to chromium acid. however. San Diego.800 Nm3/h) is operated with a transmembrane pressure difference of 60 bars. consisting of 8 hollow fiber modules (total feed capacity 3. It should be mentioned that more recently synthetic membranes have been developed (i. Wilmington or the PBIL of Teijin Ltd. Asymmetric phase-inversion membranes like the membranes employed in reverse osmosis are difficult to prepare as gas permeation is much more sensitive to micropores than RO due to the much higher diffusion coefficients of gases. the top layer of the asymmetric support structure is responsible for the separation while it is the sole duty of the coating to plug the micropores. In this case.
Ji = Qi
(PlXi -
P2 Yi)
at least as long as the conditions of the gasesare well above the critical point. On a commercial basis. the flux of a permeating component is proportional to the difference of the partial pressures at both sides of the membrane.37 shows the flow sheet for the recovery of Hz from the bleed of the synthesis loop of ammonia plants. the modules for the recovery of Hz are arranged in a “onestage-two unit” form.e. For the same reason. Gas Permeation With few exceptions. In ammonia synthesis. the bleed is utilized for heating purposes in the reforming stage. the PA 300 of UOP. Figure 6. a recycling of the unreacted components is mandatory because of its low equilibrium conversion rate. copper cyanide and zinc cyanide baths at extreme pH values. a bleed is necessary because otherwise the concentration of impurities would increase to an intolerable level. The first unit. (2) the temperature of the Watts nickel bath is 50” to 70%. still being a very common membrane material).
. the permeate can be fed to the second stage of the synthesis feed compressor.

37:
ammonia
Flow plant.
sheet for the recovery
of Hz from the synthesis loop of an
.
Feed
AMMONIA SYNTHESIS
-
I
+ Ammonia
Figure 6.
I I
i 25bar
I
-1.388
Handbook
of Industrial
Membrane Technology
5 0 Retentate 4 ’
9Sbor
Water
.

The retentate is utilized for heating pur?osesz6 Since flux and selectivity increase with increasing transmembrane pressure difference and the modules can tolerate pressure differences of about 166 bars.
phase in-
(2) the feed contains small amounts of HCI. no compressors had to be installed. spiral wound modules with asymmetric cellulose acetate membranes are employed.3927 shows a two-stage cascade for the purification and dehydration of sour gases. the permeate leaving at 25 bars as an additional feed to the first stage of the compressor.27 This process seems to be interesting for two reasons:
(I)
spiral-wound modules are employed with dry asymmetric version (cellulose-acetate) membranes. In this case. The reason is that the Hz recovery system has been added to an existing plant and that the first stage of the synthesis compressor would not accept the permeate flux of both units.
ISOEiJTANE PRODUCT GAS
Figure6.
. the tail gas leaving the conventional fractionating system contains about 70% HZ.38 shows another example of Hz recovery. that in this case as in all other cases discussed here. Again. Figure 6. It should be noted. The membrane unit placed behind the fractionating system recovers about 90% with a Purity of >96%. mainly removing CO2 and H2S.Process Design and Optimization
389
The retentate of the first unit is fed to the second unit (and for this reason. here from the tail gas of a high pressure synthesis (UOP Butamer process).38: 27 ess.
Hydrogen
recovery
from the tail gas of the UOP Butamer
proc-
Figure 6. the reader might question why the first unit is operated with a transmembrane Pressure difference of only 60 bars. This is the main reason why these applications show excellent payback times. it cannot be considered a two stage cascade).

5
16872 753 18. This can be achieved by applying either a vacuum at the permeate side or by sweeping the permeate side with a carrier like air or Hz0 vapor (Figure 6. Phase change occurs because the partial pressure of the permeating components is lower than the corresponding saturation pressure.390
Handbook of Industrial Membrane Technology
TO FEED @
AMINE
UNIT
-
8
N
TO SULFUR
RECOVERY UNIT
VOLUME FLUX I Nm3/h I MOLAR [kmol FLUX /h I
17110 763 66.27
Pervaporation Pervaporation differs from all other membrane processes because of the phase change of the permeate.5
12625 563 65.2
12712 566 17.39:
Purification of sour gas by a two-stage stripping cascade.5
29830 1331 66.
.35
PRESSURE lbarl
Figure 6.2
L130 10L 1. Masstransport across the membrane is not achieved by elevated pressures at the feed side as in RO and in gas permeation but by lowering the activity of the permeating components at the permeate side.40).

.

dh = 4 mm). the temperature decrease along the membrane. Pervaporation will always be a process which is relatively expensive compared to other membrane processes for two reasons: (1) The modules must be designed for a low pressure drop at the permeate side despite the increasing volume of the permeate due to the phase change since the principle of pervaporation is very sensitive to such pressure losses.43 shows. The necessary heat-flux . According to Figure 6. it will be the far better solution to achieve a quasi-isothermal operation by a combination of modules and heat exchangers connected in series. for the pervaporation of benzene/cyclohexane and for symmetric PE-membranes. ” Quite often the heat-flux is small and. the correlation between measurements and calculations (straight line) based on
is fairly good.41 shows that the selectivity decreases with increasing pressure at the permeate side. In any case. The dotted lines in Figure 6. Especially in pervaporation of water mixtures with water as the preferably permeating component. (2) The process requires heat transfer surfaces because the heat of evaporation necessary for the phase change of the permeate must be supplied to the process (and must be rejected by condensation of the permeate). heat transfer between the bulk of the liquid and the membrane surface (permeate-side) can become the rate-controlling step.42.392
Handbook of Industrial Membrane Technology
Figure 6. Such a design is expensive. For a ratio of p/p. the separation characteristic is identical to the thermodynamic equilibrium curve (for 19= constant). Figure 6.42 shows the temperature drop orthogonally to the membrane as a function of heat flux for water/isopropanol flowing in a rectangular channel (laminar flow. Reh = 498. for this reason.43 are valid for external heating of the channel wall opposite to the membrane. = 1 finally. In most cases. therefore. the temperature decrease in the direction of flow has to be taken into account. Due to the laminar flow. The temperature distribution has been calculated numerically assuming laminar flow in a rectangular channel with membranes at only one side.I = Z Ji
(Ahvi
+ c
piv
A91
is usually drawn from the feed and will inevitably lead to temperature gradientsorthogonally to the membrane as well as in the dirction of flow. the temperature drop orthogonally to the membrane can be neglected. heat transfer orthogonally to the membrane is only achieved by heat conduction and. the channel height should be small.
. Figure 6.

43:
Temperature
distribution
of the feed along a membrane
channel.42: Temperature bulk of feed and permeate.S.3.=O. :o % :180Wlm2
I
/
-q
’
’ ’
10’
I I
I
I I II
lo2 L
I
mm
I
I
IO3
Channel Length
Figure
6.Process Design and Optimization
5‘C I-
393
3-
-Membrane Mcmbrone Srm I
2-
l-
o’
I 1
I
Heat
I
I
4
ICAl I 1
1
0 I I
I I 1
1500
2000 FLUX
2500
3000
W/m2
Figure 6.
.
difference
orthogonally
to the membrane
between
I
BenzenelCyclohexane PE -Membranes Rear= SO.=L --ile.wi.

In most cases. the calculation was aimed at minimizing the sum of all permeate flows in the cascade. Here.44 clearly indicates that for a decreasing selectivity of the module the specific process costs will increase for two reasons: (1) the number of stages increases and (2) the membrane area increases because the internal permeate (reflux) flows increase.
10
a
6. cascades are designed for either constant reflux. Figure 6. Figure 6.44: Optimisation of a reflux cascade for pervaporation: influence of selectivity on total permeate flow and number of stages.44 shows some results of a numerical optimization of pervaporation cascades. the permeate flows are of prime importance and.32 Assumptions: (I) the selectivity is constant and equal for all stages. a reflux-cascade has to be designed.
Number of Stages
a
10
20 Sn
Selectlvlly
Figure 6. a different optimization procedure has been chosen: with respect to costs. There already exist applications of pervaporation on a technical scale.394
Handbook of Industrial Membrane Technology
The major field of application for pervaporation will be either the separation of organic components with almost identical boiling characteristics or azeotropic mixtures. (2) the fluxes are small enough to neglect concentration polarization. therefore. In this case. the desired product quality cannot be achieved in a single step-either a combination of different processes or a “multi-stage” process will be necessary.
Mostly. A
. or constant cut rate or as a so-called “ideal cascade”.

The fractionating of a 50% benzene-cyclohexane mixture into products of 98% purity is normally achieved by extractive distillation with furfural as a carrier.46 is needed.45: Permeate
I
Vacuum
Hybrid-plant for production of pure alcohol. This leads directly to the necessary number of heat exchangers and membrane units per stage and to the necessary heat transfer area. at best a three stage cascade with the refluxes indicated in Figure 6.
Although the high separation potential of permeation has been demonstrated for a number of systems in laboratory experiments.5% by only one pervaporation stage. In Figure 6.46.
1
r--l
I
I
Ethanol
99.Process Design and Optimization
395
hybrid process (Figure 6.45) for the production of pure alcohol. proved superior to the conventional approach (extractive distillation) with respect to specific energy consumption (1 kg/kg ing to Figure 6.a r
I
__ _
%
I
waste Feed Figure 6. In a numerical calculation.30 pervaporation seems to be economical only in cases where high product purities are required-and in combination with a conventional separation process.5 K). This statement is based at present only on a detailed analysis of the separation of benzene/cyclohexane and corroborated by the abovementioned production of pure alcohol.45. this temperature decrease is set equal for each module (7.33 It should be noticed. the temperature decrease “along” each module is a free parameter. where each stage consists of a combination of membrane-modules and heat exchangers. but there can be no doubt that the results are valid for many other systems as well. fluxes will vary markedly with the feed concentration of the
. The flow-sheet indicates a major difficulty of the process: if the membrane selectivity is high. that for this range of concentrations membranes for which water is the preferable permeating component must be used. the alcohol concentration is increased from about 80% to 99. If pervaporation is to be employed instead. combining distillation and pervaporation.

COOLING
WATER
-‘t’T
HEATING
Figure 6. compared to pure extractive distillation.
A hybrid process according to Figure 6. benzene will “enrich” in this section).5% are required. Permeate fluxes might differ by a factor of 10 between the first and the last stage if the same membrane is employed throughout the cascade. Compared to extractive distillation. Especially noted should be the high costs for condensation-a direct consequence of the low pressure (150 mbar) at the permeate side. even for very conservative assumptions. (Entrance temperature of the cooling water was assumed to be 2O’C. In this case.5 K per membrane unit).4). The calculations are based on the following assumptions:
. a pervaporation cascade will not be competitive (Table 6. investment costs as well as operating costs are much higher.47 with a one-state pervaporation for final purification of the top product of. the effectiveness of the carrier is rapidly decreasing in the section above the feed tray for the carrier and as a consequence. the extractive distillation shows a cost advantage of 20%.46: Flow sheet of an optimized cascade for the separation of a 50% benzene-cyclohexane mixture.396
Handbook of Industrial Membrane Technology
preferably permeating component. (a19 = 7. product quality 98% benzene and cyclohexane.) The situation is different if product purities of about 99. the energy consumption of the extractive distillation is very high because of the high reflux ratio (in the extractive distillation tower.

number
is being developed appearing
based on fermentation
of papers and patents the existence dealing with enzyme
in this field
indicate
of a new branch
of biotechnology. the perwell beyond both
and at the same time regulate their metabolic in a specific efficient leading to a substantial is extremely catalysts.
reduction
and selective.
of biological
of a living organism living
known as “metaboto cellular cooperate. in enzymatic
separation.” and the development Referring their lism”.
in colonies state.
of their applications literature
to specific suffice
‘r2 for a detailed
presentation
of enzymes
it to say that enzymes reactions confined within
are globular cells.
that prospects of in universisuch as: to pro-
possible applications
has prompted
an enormous
ties and industries (a)
leading to a series of related developments. interest activation energy. in recent years known as “enpurification and catain laborato more
scale and industrial approaches. enzymes or whole cells on
Engineering
of techniques
to immobilize
or in solid supports. 401
. (b) (c) Improved enzyme purification techniques. catalytic
production.
formances
of synthetic
It is not surprising. they properties. reactors. often
proteins which bound
lyze the complex membranes. a role similar to that of synof reaction therefore. They catalyze Although thetic their their catalysts.
and synergistically
contributing
to the energy balance of living cells reaction.
are usually Enzymes specific enzymes action
are organized reactions play.
Gabriele lorio and Gerard0 Catapano
INTRODUCTION
A growing tory traditional The clearly interest in the use of purified applications generally enzymes as biocatalysts as an alternative processes.
zyme engineering. of the techniques to induce
Enhancement
microorganisms
duce selected enzymes.7
Enzyme Membrane Reactors and Membrane Fermentors
Enrico Drioli.

immobilized (a) (b) (c) (d) (e)
Opportunity Cost cut-off
to design processes in a more rational in enzyme plants.
More compact Operating
costs cut-off.
per unit time.
on a surface.3 Several advantages can derive from the use of cells. with small amounts
High productivity of side-products.
of
techniques
toward
the
solid
phase synthesis
of
Enzymes cals. immobilization
has therefore and procedures techniques
been performed
in order to optimize of enzyme
in view of the development include:
Confinement Encapsulation Adsorption
in a gel. be of extreme in terms preservation. in maintaining above problems standard could product quality. interest it implies of energy
The use of enzymes individual applications safety. which or microbial
and can be re-
and continuously enzymes
used”. per unit time per equipment. consumption.
by microorganisms fermentation under etc. and reuse of enzymes or cellular
and expensive
recovery
microorganisms. using “immobilized properties en-
Difficulties of the
be overcome catalytic
or in other region
words
“enzymes retain
confined their
or compartmentalized
in a well
of space. of chemically can therefore and materials stirred
of pharmaceutiside-product
reactions
perform
the synthesis as biocatalysts mainly prevention reasons: costs. The availability
such as fungi or bacteria of almost pure enzymes to limit active
have been enables one forwhich for
used for years in batch to carry out specific to otherwise
plants for the production mild conditions..402
Handbook (d)
of Industrial
Membrane for
Technology enzyme usage in continuous flow
Development reactors. with low yields. ever limited (a) (b) (c)
for the advantages in batch
pollution
The traditional
use of enzymes
and similar reactors is how-
for the following
High enzyme
purification
Low productivity Difficult
per reactor.
A lot of research work immobilization Traditional (a) (b) (c) techniques reactor engineering. defined peatedly Product pollution. such as: way. mation would and
produced foods.
.
compounds
require extremely
long reaction
time.
of techniques
(e)
Development peptides. (d) (e) Most zymes”. beverages. consumption. in a membrane shell.

in nature a continuous uptake of substrate and release of product without loss of biocatalysts is not achieved by carrier fixation but by means of cellular membranes. in this sense. such as cofactor-requiring mono. Efficient immobilized enzyme reactor systems for technical applications can therefore be established replacing cellular membranes by ultrafiltration or reverse osmosis synthetic membranes. This limitation. to further developments of immobilized enzyme reactors for industrial processes.and multi-enzyme systems. (e) Copolymerization with a proteic carrier. which are the main drawbacks of the aforesaid enzyme membrane reactors. can nevertheless be used to form. On the other hand. more than seventy enzymes have already been immobilized. Enzymatic systems in which membranes are simply used as separation media and not as catalyst carriers are traditionally called “enzyme membrane reactors” (EMR). systems where product concentrations are kept low and a continuous product removal occur are particularly interesting when dealing with product-inhibited enzymes. Fluid dynamic conditions in some of these reactors make them especially suitable for enzymatic systems for which a homogeneous catalyst distribution is particularly important. magnetic fields or similar apparatus. but only ten of them are used in most existing applications. in fact.’ Ultrafiltration membranes can be used to retain enzymes in the reaction vessel due to size difference between the usual high molecular weight enzymes and most of the low molecular weight substrates and products. Concentration polarization phenomena severely affect the performance of such reactors so that it is necessary to control the polarization layer onto membrane pressurized side by means of reactor fluid dynamics or design tricks. and the development of practical equipmenL4 Membrane science can strongly contribute. either in dynamic or in static conditions. Depending on enzyme kinetics. requires that research efforts be directed towards the realization of more efficient immobilization techniques characterized by low carrier costs. different reactor configurations exist which are able to optimize system yield. a homogeneous continuous catalysis can thus be achieved. Some membrane reactor configuration do not allow a homogeneous distribution of catalyst in the vessel so that the reactor acts in a more heterogeneous way. and can then be recovered by means of sedimentors. This well assessed technology seems.2741 It is even possible to establish more than one enzyme layer and no coupling agent is needed to carry out the immobilization.Enzyme Membrane
Reactors and Membrane
Fermentors
403
(d) Covalent binding to an insoluble support. very promising for the design of compact and flexible apparatus and particularly useful when a separation step is to be coupled to the chemical reaction. a leading country in the development of enzymatic engineering. and the activated transport through the cellular wall by a forced flow across the membrane.
In Japan. for instance see References 8-26.6 Insoluble carrier-fixed enzymes can be continuously used in stirred vessels or tubular reactors without catalyst loss. Concentration polarization phenomena. a gel layer of enzyme proteins on a membrane. compared to the great potential of immobilized enzyme systems. Due to high protein
.’ A number of applications of membrane processes to biotechnology are already in operation.

we will refer to five main reactor Enzyme membrane reactors. where enzymes are immobi-
lized in a proteic (c) (d) (e) Membrane Membrane Membrane
gel layer. where
enzymes
continuously
flushed along membranes. acts mainly
can only for entrap-
to and from the biocatalyst.404
Handbook
of Industrial
Membrane surface. viable cells are grown for a wide range of applications.70q3 and used in membrane will range from fermentors systems prolevel. L* L
of the fiber.
in continuous-flow
systems. where
Last but not least. 42-53 The enzymatic although
thus be spread In certain the membrane mass transfer Enzymes in order from elution.
. accessible to substrates and products preventing can still or simply membranes. In the
sections. is not suitable membranes.
LIST
OF SYMBOLS The following symbols are used throughout Membrane surface Inner radius area. zone in the by diffusive caused
it defines a reactive mainly or inhibition
vessel. the catalyst be absorbed covalently from pollution within by other species in soIution. applications. Asymmetric for enzymes. dynamically reactors.54-6g symmetric bound
possibly
macroporous either
membranes them or can
to establish
high catalyst
concentrations. or ionically membrane
cross-linked times. (b) Dynamic enzyme gel layer reactors. through either the unenzymes the can
A biocatalyst
of asymmetric
cells. into the supporting In these cases.
segregated enzyme bound enzymes fermentors.
to prevent
to symmetric
asymmetric be achieved
In spite of short residence
high conversions
in most kinds of enzyme
reactors. lo6 Reduced possibility tem disadvantages.“-g8 Our and from enzymes following (a) analysis of enzyme membrane
reactors
posed at bench
scale to equipment enzymes confined
already are bound
in operation
at the industrial to apparatus
systems where are simply
to membranes
in well-defined
regions of the reaction configurations: are
vessel. the size of the biocatalyst sponge layer of asymmetric as a selective barrier. usually the shell. although spongy
still suspended.
are effectively substrates
immobilized and products
within activity
macroporous diffuse ment
part of the membranes.
Technology enzyme stability distributed can be improved in the reacting limitations over solusys-
concentration that
on the membrane using enzymes catalytic
in systems of
homogeneously
tion.
over a large surface. skinned or whole surface
efficiency pathways fiber
due to mass transport in the enzyme
and the
preferential hollow
gel layer can be typical
membranes suspension membranes
can also be used as selective supports can in fact so that be forced biocatalysts.
or statically
formed. this chapter.

T Downstream dimensionless time constant (time lag) Downstream dimensionless time constant (segregated region)
21 = T1 v* /D
72 = T2 v* /D
0
=(-
K E 6* D Sf K E A2
)l/*
Thiele modulus
V'=( 6 = 1'2
)w
Heterogeneous gel model Thiele modulus
D 72(O) Sf 8' Dimensionless substrate concentration within the gel layer -22 2
x = S /-f*(O) Sf
Nz)
. M L-l T‘* Dimensionless reaction rate. and the existence of diffusional resistances limits this approach with low activity enzymes.Enzyme
Membrane
Reactorsand
Membrane
Fermentors
499
Oncotic pressure. A number of drawbacks affected the use of immobilized enzymatic preparations. the attention of applied research has been focused on the engineering of systems based on immobilized biocatalysts. M L-l T-* Shell side pressure. with macromolecular substrates and in general with enzymes whose cata-
. Enzymes involved in this development were enzymes catalyzing simple reactions that normally require no cofactors.e
0 I
II* e du
Dawson's integral
o =(d-a)
DL/(a Ds)
Ratio of diffusion rate in the bulk phase to that in the
porous catalyst
ENZYMEMEMBRANE
REACTORS (EMR)
During the last ten years enzyme technology has moved mainly towards the development of new immobilization techniques and the improvement of those already existing. An often dramatic reduction of initial enzyme activity due to the binding process. In turn. P = Q1 7s P = Qo
7 =
for first order kinetics for zero order kinetics
V/Q
Reactor time constant.

mixing of substrates and catalysts is not fully accomplished. Biotechnological processes based on immobilized biocatalysts have not been successful with coenzyme-dependent reactions. An EMR is a reactor system where ultrafiltration or dialysis membranes with a suitable molecular weight cut-off are used in order to keep larger components in the reactor vessel. The direct and deep contact between substrates and biocatalysts limits diffusional resistances. LC: level controller.g. i.1).
@N2
Membrane-l
Figure 7. thus leaving the reactor as permeate. Instead of being immobilized. enzymes and/or macromolecular substrates.1: Experimental set-up of an enzyme membrane reactor. while no activity losses due to fixation to the carrier occur.. and concentration polarization phenomena strongly limit reactor performance. together with easy recovery of deactivated enzymes and replacement with fresh catalysts are achieved. Moreover.” P: peristaltic pump. on the other hand. In this set-up. So far. typical advantages of immobilized preparations.
. Performance of dead-end units is largely affected by the flow dynamics of the system. thus maximizing the activity of the biocatalyst. enzymes or cells can be confined in the vessel of dialytic or ultrafiltration cells. products and/or inhibitors. Depending on the flow dynamics of the reaction vessel. the existence of a polarization layer above the membrane surface significantly reduces permeate flow rates. e. are allowed to pass freely through the membrane. inhibitors are moreover continuously removed from the reaction vessel.e. enzymatic conversion is achieved mostly inside a highly concentrated polarization layer giving lower conversions as compared to those obtained with the same amount of enzyme uniformly distributed within the reactor. in a configuration that is commonly called “enzyme membrane reactor” (EMR). both dead-end and CST UF cells with flat membranes have been proposed as enzymatic reactors (Figure 7. in fact. while low-molecular-weight molecules. On the one hand.410
Handbook of Industrial Membrane Technology
lytic behavior is strongly dependent on a good mixing of catalyst and substrate. it is possible to distinguish between reactors equipped with flat UF membranes or hollow fibers. removal of deactivated enzymes or cells from the solid support often raises serious problems and sometimes is not convenient.

X)1 KIM = KEV/Q (3)
R being expressed as K E S/(K’M + S). it must be outlined that the reaction term is the course of the reaction rate relative to the substrate conversion of the key component. it furnishes a rapid graphic procedure to obtain rough estimates of kinetic parameters. biocatalyst losses can even occur due to the wrong choice of membrane molecular weight cut-off. True reactor operating point at steady state is given by the intersection point of the straight line representing mass balance and the curve of the reaction term.Pf)/r = R(S. the slope of the convective term is reciprocally proportional to the mean residence time.st10r99As can be seen from Figure 7. If multienzymatic systems are used as the biocatalyst. combining Equations 1 and 2 with the rate equation eventually leads to an implicit Equation:9 (P .Pf)/r = R(P) (4)
Product concentration in the permeate can then be determined either with numerical procedures or with graphic techniques.P) (2)
Assuming that substrate conversion obeys the simple Michaelis-Menten model. The equation in this form is a useful tool to estimate parameters of reaction kinetics. substrate steady state mass balance reduces to: XSf + [X/(1 . and the latter is equal to its value in the permeate.Enzyme Membrane Reactors and Membrane Fermentors
411
CST reactors have been more widely adopted. When enzymes or cells are compartmentalized in UF cells. Instead of performing nonlinear parameter estimation procedures.X) can be plotted. the functional dependence of XSf on X/(1 . tx. When more complex kinetics are involved. Hence. for cu-ketoisocaproate conversion to L-leucine. that is the time at which enzyme activity is reduced to half its initial
.P) (I)
Product steady state mass balance similarly appears as: (P . and partly due to the easy modeling of enzyme kinetic behavior.2. It is conventional to measure the enzyme stability in terms of its halflife. partly due to the possibility of concentration polarization control. so that the substrate rate of conversion is dependent on both substrate and product concentrations. comprehensive mathematical descriptions of the kinetic behavior of enzymes located in CSTR UF units are reported8t9r’e. Physiochemical changes in enzyme structure. Assuming complete mixing within the reactor so that enzyme and substrate concentration in the reactor vessel are uniform. substrate mass balance in molar form generally looks like: (Sf-S)/T = R(S. In the literature.as far as low-molecular-weight substrates are concerned. thermal denaturation and microbial contamination cause enzyme activity to continuously decrease with time.” Enzyme activity is usually not constant with time. The plot should look like a straight line whose slope is (-KIM) and whose intersection with the XSf axis is given by the point of coordinate (Vmaxr).

4). =
1.o/I(1/7
.6
01 0
I
I
I
I
I
1
10
20 t. while biocatalyst
. exp( . d
30
Figure 7. For macromolecules (like most biocatalysts). tr/. P) 0. e. for various retention times: (A) 0. low-molecular-weight products leave the reactor permeating the membrane. (0) 0.Enzyme Membrane Reactors and Membrane Fermentors
413
Figure 7. When the reacting solution is fed to the system. The extent to which concentration polarization affects reactor performance depends on the balance of rejected solutes... membrane rejection is usually very good.S(t)1 /T .8 For 057.Kd) (Sf .417 days. can be expressed as:
tr/. accumulation due to membrane rejection and back diffusion to the bulk phase... if membrane properties are carefully chosen.520 days. whereas enzymes are partially or totally rejected. and eventually on the flow dynamics of the reacting vessel.t/r) contribution can be neglected so that enzyme half life time.Kd) and the line intersects the vertical axis at the point of coordinate (In V. enzymes tend to accumulate in a thin layer immediately upstream from the membrane causing polarization phenomena to occur.3: Experimental and (solid lines) theoretical plots of substrate concentration (S) as a function of time. enzymes or cells.695 days8
Estimate of the deactivation constant can be graphically carried out from Equation 7 plotting In [[Sf .).g.dS(t)/dtl vs t (Figure 7. Then. A straight line is thus obtained whose slope is given by ( ..3 shows good agreement of experimental data and theoretical predictions for benzaldehyde conversion by Rhodotorula Mucilaginosa.6
InjV.So)/2jl/Kd
(9)
0.

integration of Equation 10 allows one to express the ratio of enzyme concentration in the bulk solution in the presence of concentration polarization.6 0 10 14
Figure 7. to its
. (C) 0.C. x + m E = Es.D.417 days.5 shows a schematic of the situation.dE/dx (10)
under the B.414
Handbook of Industrial Membrane Technology
0. Applying the thin film theory to a region immediately upstream from the membrane results in the following steady state mass balance equation:‘w~lO1 J E = .695 days. Figure 7. Under the assumption of totally rejected enzyme macromolecules. Straight lines are approximations of experimental data Ito a straight line calculated from Equation 7.
0.4: Values of Vmaxo and Kd for various retention times: (A) 0.8
20
back diffusion towards the bulk phase is extremely slow. The effect of concentration polarization on such reactor performances can then be significant. (B) 0.8.520 days.

more generally Ks. Doubling flow rate from 4 to 8 ml/min results in an B-fold decrease of fractional bulk concentration of soluble enzyme. Changes in enzyme bulk concentration may induce a large reduction in reaction rate within the membrane reactor. from 97% to 12% for the cellobiose/cellobiase system.Enzyme Membrane Reactors and Membrane Fermentors
415
1
Figure 7. and hence applied pressure.7 and 7.6 shows how dramatic the change of enzyme bulk concentration can be. D/6. with changes in permeate flow rate. = I/[(1 -a)
J*A
Schematic diagram of enzyme membrane reactor.8 give evidence of the existence of enzyme deactivation phenomena which can be misleading in the evaluation of reactor performances. relative to the overall mass transfer.5: ES/E. and/or of the thickness of the boundary layer close to the membrane surface in accordance with Equation 12.’
value in the absence of concentration polarization phenomena as:s + (a/P) [expW)-
111
(11)
Enzyme concentration at the membrane-solution interface can thus be related to enzyme concentration in the bulk phase by the following equation: EVV = Es expU3 (12)
where the dimensionless parameter fl is a measure of convective mass transfer.
.“* Figure 7. J. Figures 7.

Enzyme Membrane Reactors and Membrane Fermentors
417
A
A
&A
A
16 Time. in a stirred UF unit proteic coiled macromolecules in the polarization layer are subjected to a high shear field due to stirring. Overall enzyme deactivation is moreover accelerated by the exchange of deactivated enzymes in the polarization layer with fresh enzymes in the bulk solution via a coupled diffusion-convection mechanism.05 ml/min. as concentration polarization occurs. Boundary layer thickness under such conditions cannot be easily estimated from the conversion at various permeate flow rates. due to the consistency of product concentration in the bulk phase of the reactor and in the permeate. Beyond the critical value.’ Correspondingly.21 mg/ml). being
. and correspondingly outlet product concentration is constant. dotted line is expected level for ideal operation (1. Experimental evidence suggests that the contribution of polarized enzymes to overall conversion can be negligible. therefore.’
Under given stirring conditions.hr
24
32
Figure 7. At flow rates lower than the critical value.8). an accelerated deactivation of enzyme activity is superimposed on reactor performance thus hindering the attainment of steady state conditions. can be attributed to the localization of large amounts of enzymes in a region immediately upstream from the membrane.8: Concentration of glucose in permeate versus reaction time: (0) at flow rate of IO.3 ml/min. Irreversible enzyme deactivation is confirmed by the discrepancy between the expected values of conversion and experimental data as applied pressure is slowed down from hypercritical to subcritical values (Figure 7. concentration polarization phenomena promote the localization of a large fraction of enzymes near the membrane surface. As suggested by Tsao et al.8 mg/ml). solid line is expected level for ideal operation (0. the reactor always attains steady operation conditions. seldom leading to an enzymatic gel formation. which can result in protein unfolding and consequent denaturation. Enzyme deactivation. (a) at flow rate of 8.9 a continuous conversion decrease. a critical value of flow rate and hence of applied pressure does exist.

more marked for lower residence times. concentration polarization phenomena affect the EMR performance more severely. Under either operational mode permeate flow rate continuously decreases with time. or when inhibiting species are assumed to exist in the feed stream. The stirring features of CST EMRs moreover assures that substrates and/or inhibitors within the reactor vessel are maintained at the lowest possible concentration level. By selecting membranes with an appropriate molecular weight cut-off.. i. both enzyme and substrate are retained in an EMR in touch with each other. Initial sudden increase in product concentration in the permeate occurs earlier at lower residence times. Figure 7. When EMRs are operated continuously. Figure 7.g. Such reactor configuration is then extremely useful when substrate inhibited reaction patterns are involved. and it may happen that. dextrins in starch hydrolysis. When macromolecular substrates are involved in the transformation under study. as shown in Figure 7. concentration polarization phenomena play a dominant role.a As far as the first two operational modes are concerned. eventually approaching steady state conditions in terms of both permeate flow rate and substrate conversion.16j’7 In addition.9 shows typical curves of product concentration as a function of time for the hydrolysis of cellulose according to a diafiltration operational strategy.10. leads to meaningful changes in reactor performance as compared to the diafiltration mode.11. Operating the reactor in a semi-continuous condition and adding substrate so as to keep its bulk concentration constant. followed by a slow decay. feeding a slurry of substrate macromolecules to the reactor. The presence within the reaction vessel of contaminants or intermediate products. higher flow rate. Diffusion limitations of macromolecular substrates hamper the use of immobilized enzymes in the hydrolysis of high-molecular-weight substrates.. A peak is always observed in the first period of operation. after a sudden increase in the early stages. Diafiltration. It is to be noted that deactivation phenomena are less prominent for enzymes stable to shear stresses. Soluble enzymes can then act directly on substrate macromolecules without diffusion limitations and steric hindrance imposed by enzyme fixation to a solid support. e.” semicontinuous’ and continuous operational modes13-15 have been proposed and applied for saccharification of cellulose11~*2~14~15 and protein hydro1ysis. the filtration rate continuously decreases with time.418
Handbook of Industrial Membrane Technology
superimposed on the enzyme concentration decrease in the bulk phase. can be misleading. and hydrolysis products and/or inhibitors are continuously removed from the system. The addition of enzymes capable of hydrolyzing such foulants to low-molecular-weight compounds usually improves reactor performance. antifouling procedures including suitable feed pretreatment or procedures to reduce concentra-
.e. Under such conditions. at different space velocities.r3 which are not fully hydrolyzed by the enzymatic system under study can lead to membrane foulingI or to the formation of a gel layer at the membrane surface. substrate conversion does not attain steady state conditions. Such findings confirm that enzymatic hydrolysis of cellulose is strongly product inhibited and that suitable flow dynamic conditions greatly improve reactor performances.

.”
tion polarization can be used.e. J.9: Concentration of soluble sugars (cellobiose and glucose) in permeate from an enzyme membrane reactor as a function of time for various space velocities (S).10: Conversion of cellulose to soluble sugars as a function various space velocities (S). diminish concentration polarization. as discussed earlier. albeit at the expense of partial enzyme deactivation.“’ Figure 7.”
Conversion(%) 60 I
4
8
12
16
20
24 of time for
Figure 7. lo1 Equation 12 demonstrates the dependence of enzyme concentration at the upstream membrane surface on the flow dy_ namics of the reaction vessel. i.
. operational conditions improving mass transfer in the bulk phase. i.12” shows how high stirring speeds usually improve filtering performance of a CST flat membrane reactor. KS or D/6.. Due to the monotonicity of the exponential function. at given applied pressure.e.Enzyme Membrane
Reactors and Membrane
Fermentors
419
L
4 8 12 16 20 24 28 32 36 40 44
l
4a
TimeIhr)
Figure 7.

in a stirred cell system. Flushing enzyme/substrate solution from a reservoir to an HF module in a semiclosed loop at a high linear velocity reduces the extent of concentration polarization avoiding the formation of a secondary enzymatic gel membrane.t.13. and that recirculation flow rate is much larger than permeate flux. Assuming biocatalyst kinetics are described by the simple Michaelis Menten model.Enzyme Membrane Reactors and Membrane Fermentors
421
Reactor configurations other than dead end or CSTR can offer improvements.= 100 . but also to the increase of the surface to price ratio. min
Figure 7. a steady state is also attained. the performance of a thin channel flat membrane reactor for the hydrolysis of sweet potato by means of a-amylase is reported in terms of filtration rate and product concentration vs time plots.____-““-3
. in Provided that the volume of spite of product flux through the membrane.r3
32.11 shows how great the improvements due only to the change of configuration can be.31& gm _____
r/(
A.13: Effect of adding crystalline sweet potato /3-amylase (87 pg protein/ml reaction mixture) on the filtration rate (ml/h) and on the reducing sugar content (measured as maltose) in the filtrate at SO%. f! a 0 200 Time . ‘s Reactor productivity as a function of time can be related to flux and enzyme concentration according to the following equation: Pr = x(t) SfJ(t)/E V (13)
300 it! sii y 200 s In .0
160 Se-l% w/v V =550 ml
E =(1066gm -
t
//
nn
a
E =0. Capillary membranes are characterized by a favorable surface to volume ratio. A comparison to Figure 7. . the substrate mass balance can be expressed according to Equation 3. The use of hollow fiber ultrafiltration modules strongly improves both reactor performance and economics. In Figure 7. 101t102 the UF module is small relative to the total volume.I
400
.
_*. system kinetic behavior can be modeled in terms of a CST reactor. advantages from this feature are not only related to lower overall plant size. Enzyme added (0). 15 psi. At an appreciably high flow rate and product concentration level. no enzyme added (o).

Enzyme and substrate concentration appear to affect reactor capacity to a large extent (Figures 7. maximum productivity. as well as more labor.10 shows how conversion rate is improved by continuous removal of inhibiting products. that is increasing reactor volume or lowering fluxes.16) especially at lower enzyme concentrations. Comparisons of EMRs operated in the diafiltration mode” and in the continuous mode” to batch reactor performances are reported in Figure 7.e. for soy protein hydrolysis performed by Pronase enzyme.14 shows how reactor productivity (g product/g enzyme) is affected by flux and enzyme concentration. Figure 7. Figure 7. g product/g enzyme/min) is affected by operational parameters. enzymes are charged at the beginning of the process. batch processes require added expenses for enzyme inactivation and product purification. the enzymes must be replaced at the end of each cycle.
E =Q066gm
_
E =Ct31C gm _____
200 Time.314 g: (0) = 9 ml/min. results in higher conversions at the expense of lower capacities (Figures 7. min
400
Figure 7.15-7.19 (reactor continuously operated) shows that productivity levels are higher the longer the reactor is operated before shutting down for cleaning.”
. In batch reactors.18 show how reactor capacity (i..15 and 7. In the continuous process. and the system is kept working until enzyme activity drops down to a given limit established on the base of enzyme half life and plant economics.18).19.17 and 7. Increasing residence time. E = weight of enzyme added to reaction vessel (flux at E = 0. Maximum productivity can be obtained by operating the reactor at the highest flux and the lowest enzyme concentration. Figures 7. (A) = 16 ml/min).10 and 7. recharging or backflushing. the greater is the productivity and the larger the difference between continuous and batch system performance.422
Handbook of Industrial Membrane Technology
Figure 7.14: Effect of flux and enzyme concentration on UF reactor productivity. Effects of permeate flux and reactor volume on conversion are related to each other by means of substrate residence time in the reaction vessel. Hence the longer the reactor is operated. Moreover.

‘0~20~21 Coenzymes.17: Capacity and conversion of a UF reactor as a function of flow rate through the system (flux). like NAD’ or NADP’.18:
1000
2mo
3ooo
Effect of reactor volume on UF reactor capacity and conversion. Enzyme membrane reactors offer a suitable reaction environment provided that coenzymes are retained in the reaction system. able to oxidize them. to another. usually have a long term effect on enzyme activity only if they can move from one enzyme. in loop kinetics.Handbook of Industrial Membrane Technology
2. Such compounds are in fact quite expensive.‘g
The real potential of membrane reactors becomes evident with coenzyme dependent enzymatic systems.”
Conversion
0
Figure 7. Reverse osmosis (RO) membranes could be helpful in retaining native
. Continuous homogenous catalysis is then a prerequisite to achieving high reaction yields.4
0
10
20 Flux (ml/min)
30
Figure 7. able to reduce them. which limits the use of coenzyme dependent enzymes.

thus limiting overall fluxes. half life. the presence of product in the permeate implies coenzyme loss due to the usually small differences of their molecular weight. Volume replacement for the continuous UF reactor calculated as J t/V. If a suitable pair of enzymes is charged into the reactor together with the macromolecularized coenzymes.‘0~‘03~104 Choosing the membrane molecular weight cut-off so as to reject the soluble polymers results in retaining the coenzymes within the reaction vessel. Patents already exist such as for amino acid production from a-ketoacids. Laboratory and plant scale experiments”t*’ show the importance of a careful choice of the enzymatic system. Moreover. some procedures were proposed in order to increase coenzyme molecular weight by covalently binding the coenzyme to a highmolecular-weight soluble polymer such as polyethyleneglycol ordextran.** which confirm the potential of scale-up for EMR plants. permeability of such membranes towards water is poor as compared to UF membranes.19: Comparison of productivity of a continuous UF reactor and a batch reactor for the Promine D-Pronase system at pH 8. In the early 197Os. thus reducing overall costs.0.19
coenzymes in a given vessel on the basis of their size. EMRs have a number of advantages over other immobilized preparations.
2
8 J f/v
10
Figure 7. Unfortunately. the latter are continuously regenerated and long term operation is assured. resolution. derivatization and racemization steps can be avoided.Enzyme Membrane Reactors and Membrane Fermentors
425
‘-“““““““““‘-‘T’
0
6 Volume Replacements. Differences in enzyme cofactor requirements. as well as the ratio of oxidizing to reducing enzymes. and optimum pH and temperature can in fact render one enzymatic system rate limiting thus reducing reactor performance. For this particular process. In summary: (a) Homogeneous catalysis
(b) No diffusion limitations
.

however.*’ Bench and large scale plants are already in operation for the preparation of N-acetyl-D. EMRs have been successfully used with macromolecular substrates.
Efficient circulation. Thus far.L derivatives from Lamino acids by means of acylase and the production of L-malic acid from fumaric acid by means of fumarase. depends strictly on enzyme concentration at the memI brane-liquid interface. In conclusion. flexibility of such systems can be considered to be limited only by the number of active compounds which can be synthesized via an enzymatic route.30 If gel formation occurs. such processes can provide an interesting technique for enzyme immobilization which takes advantage of concentration polarization phenomena. they are still in soluble form. A large membrane area and sterile filtration are also needed to assure high yields over long periods of time.426
Handbook of Industrial Membrane Technology (c) No activity lossesdue to carrier fixation
(d) Absence of enzyme fixation costs (e) (f) High activity per unit volume Sterilizability of the plant
(g) Constant productivity assured by enzyme dosage (h) Absence of contaminants (i) (j) Interchangeability of substrate/enzyme systems Use of multienzyme systems. and hence their dynamic immobilization. their settlement is partic-
. as for cellobiose hydrolysis’ and Lmalic acid production from fumaric acid. which becomes quite a problem as enzymes are added to keep conversion constant. is required to prevent concentration polarization. as for the saccharification of cellulose11-13~15~26 protein hydrolysis. conversions of up to 86% with resolution rates of up to 85% have been attained. enzymes are effectively immobilized without meaningful changes in their microenvironment. Although they are confined near the membrane surface at fairly high concentration levels. these molecules will accumulate on the active membrane surface and possibly deposit there as a thin gel layer characterized by enzymatic catalytic activity.
DYNAMIC
ENZYME GEL LAYER
REACTORS
Ultrafiltration of protein solutions is a proven unit operation”’ for obtaining purified enzymes from cell cultures.20).” and or with low molecular weight substrates. enzymes are not immobilized.31 Moreover. shear forces may deactivate enzyme molecules.*’ In the case of fumarase.29 (F’g ure 7. When the maximum enzyme concentration is lower than the gel concentration value. Under certain circumstances.*‘r** Actual gelation of enzyme proteins. If an enzyme solution is flushed under pressure through a UF membrane that completely rejects enzyme molecules. If the enzyme solution is to be flushed at high linear velocity near the membrane surface.

C. i. defined as the ratio of gel volume (in which
. x.a steady state mass balance on retained species. x is the distance from the membrane wa1l.. When unstirred batch membrane units are used as reacting vessels.31
ularly attractive due to the high enzyme and protein concentrations in the gel. the latter strongly enhancing enzyme stability. (for x = 0) and it can be expressed as:
Ew =
Nv/(ADe)
(17)
Enzyme depth of penetration.lo6 Enzyme gel layers can be built up under a number of fluid dynamic conditions.Enzyme Membrane Reactors and Membrane Fermentors
427
C
bulk
“bulk
membrane
gel-layer
Figure 7.e. Taking into account both convective and diffusive mass transfer mechanisms. enzymes.1 B.2 x=0 DedE/dx+vE=O E(x) A dx = N (15)
Upon integration of Equations 14 and 15. Unstirred and stirred batch reactors have been used as well as systems where the enzymatic solution is kept continuously flowing along semipermeable membranes until gel formation sets in. the enzyme concentration profile appears to be:30r32 E(X) = [Nv/(ADe)I
exp[-vx/(De)l
(16)
Maximum enzyme concentration occurs at the upstream membrane surface. that is: B. leads to the evaluation of their concentration as a function of the distance from the membrane surface. the enzyme mass balance equation at steady state is as follows: Ded2E/dx2 +
V
dE/dx
= 0
(14)
Proper boundary conditions can be derived assuming (a) complete enzyme rejection and (b) no enzyme loss.20: Gel layer formation ultrafiltering a protein solution through a semipermeable membrane.C.

As far as an IEMR is concerned.6 De/V (18)
Enzyme average concentration in the gel is therefore 21% of enzyme concentration at the gel-membrane interface. respectively. Comprehensive models for an immobilized enzyme batch membrane reactor (IEMR) and for a soluble enzyme batch membrane reactor (SEMR) are proposed in References 33 and 30. a steady state analysis is not useful for determining whether gel formation has occurred.21 and 7. Similar equations under different fluid dynamic conditions can be derived from Michaels’ gel formation theory”’ or from similar theories modeling concentration polarization phenomena. With the enzyme gel concentration known.10’~102~‘07~108 Unfortunately for most substances. Assuming that neither reactant nor product is rejected by the membrane and that diffusion coefficients of
.30t33 Figures 7. can be estimated to be:30 xe = 4. In some cases.428
Handbook of industrial Membrane Technology
99% of the initial amount of enzyme is contained) to membrane cross-sectional area. Ultrafiltration of protein solutions usually results in a time progressive decay of permeate flow rate. for a flat slab membrane configuration. which after a period of time attains a steady state constant value for a given transmembrane pressure. and only order of magnitude estimates can be gained from the literature. Referring to the up stream region. N. Equation 16 can be used as a design equation in order to estimate the amount of enzyme. any further transmembrane pressure increase does not enhance permeating fluxes correspondingly. Such a decay can be attributed either to gel formation. analytical closed-form solutions are available for limiting first and zero order kinetic rate equations. an accurate knowledge of gel concentration is not available. If a gel is formed.22 show typical results for the two cases-with enzymatic gel layer formation and when soluble enzymes are confined only near the membrane surface. An effective distinction between enzymes in gel form and enzymes in the soluble state can be achieved by the analysis of the transient behavior of the enzymatic reaction. the mass balance must take into account both diffusion and convection due to the usual low permeating fluxes. The steady state kinetic behavior of such an enzyme membrane reactor is only related to the kinetic behavior of enzymes in the gel or in the soluble state.33 The unstirred reacting vessel is divided into two sections-upstream and downstream of the membrane supported enzyme gel layer.“’ Protein gelation. Criteria need to be developed in order to experimentally detect enzyme gel build-up. required to build up an enzymatic gel layer of thickness x. or to an increase in osmotic pressure of the ultrafiltered solution as a result of macromolecule accumulation at the membrane-solution interface which lowers the pressure driving force. for instance. is thought to occur at concentration levels ranging from 20% to 40% by weight.3o permeate flux decay associated with gel formation is negligible thus ruling out the possibility of detecting eventual gel formation. Unless knowledge of intrinsic enzyme kinetics both in the gel and in the soluble form is available.

=
1
. The downstream region of the UF test can be modeled in terms of a first order system in series to a pure time lag.C.430
Handbook of Industrial Membrane Technology
either compound are constant and close to each other.1 states that the height of the region where substrate and product concentrations appreciably differ from their feed values is negligible compared to cell height.
Yp =
0
(21)
y/1
y
P
=o
B.2
Boundary condition B.ays ays ae x2 ac
(19)
(20)
with the following boundary conditions:
e=o c-+-
Y.
Yp (s)
e
-
Tl
s
(22)
u P
(0.25
O* + @CO*)+
al+
1
(23)
.C.
.(e)
=
I
erf “1 + 1
J 0. B. mass balance equations in dimensionless form can be written
-+-=.2 stems from the assumption that enzymatic reaction takes place only on the enzyme layer free surface.5
r2 ) --
Y.s)
(1+
‘c*
s >
Overall system response in time domain can then be evaluated from the system response in Laplace domain for first and zero order kinetics:
CL1 +
0. Dimensionless product concentration at the UF cell outlet is then related to that immediately downstream of the membrane by the system transfer function.C.5
( 1 e
_-
e*
0.

C.23: Hydrolysis of sucrose (3 x IO” mols/ml) by immobilized invertase operating in an IEMR. Immediately upstream of the membrane. Comparison between experimental data (expressed in terms of glucose.
-zero .* . enzyme macromolecules are pressure driven towards the membrane surface.32
. 0 20 40
I 60 t .(9) = (26) ( al + 1) e+m
lim
Yp(e) =
a0
e+m
When the amount of enzyme is insufficient to ensure that the enzyme concentration at the membrane surface attains the gel value. reaction takes place only in the region where enzymes are confined and whose thickness is determined by Equation 18. Once steady state conditions are attained.
order . As time goes by. As soon as the substrate stream is fed to the system. A soluble enzyme membrane reactor (SEMR) is thus defined.432
Handbook of Industrial Membrane Technology
Figure 7.) and theoretical curves (solid lines) calculated according to Equations 23 and 24.. . thus reaction rate in most of the cell volume progressively decreases due to an enzyme concentration decrease. the reaction rate increases due to an enzyme concentration increase. min
1 80
Figure 7.23 depicts a model fitting experimental data for the enzyme invertase for both kinetic orders3’ An overall steady state mass balance relates kinetic parameters of either rate equation to the ultimate value of the dimensionless product concentration yielding: ai lim Y. *
0. enzyme remixing occurs throughout the vessel and reaction then starts throughout the whole cell. Equation 17 indicates that the enzyme is confined in soluble form at a fairly high concentration in a region immediately upstream from the membrane.

/dt
(27)
Regarding region B as a lumped parameter system of volume Vb (with Vb=xe A) fed by the outlet from region A. Region A. Assuming a uniform enzyme distribution within the cell volume at system start-up. As far as region A is concerned.24: Actual and lumped parameters system. the average enzyme concentration
Actual
System
Lumped Parameters System
Figure 7. to give Es(t) = (N/Va)exp(t/ra) (28) t = 0 Ea = N/V. VA is the dominating volume portion where enzyme depletion takes place. Vs is the constant volume in which enzyme accumulation occurs. the functional dependence of enzyme concentration in region A on the reaction time is obtained upon integration of the mass balance equation: E.Enzyme Membrane Reactors and Membrane Fermentors
433
Simple mathematical modeling of this situation is reported in Reference 30 referring to the equivalent lumped parameters system in Figure 7. namely region B.= subject to the initial condition I. where a time dependent enzyme depletion occurs.24.
7a
dE.3o
. Each region is modeled in terms of a single well mixed tank.C. for a particular UF test cell this assumption can be verified simply by a tracer step response test performed on the system itself. virtually coincides with the cell as a whole except for the negligible liquid volume where enzymes accumulate. Also in this case the reacting volume is modeled as two separate regions in series to each other.

Ra. thickness of region B. The kinetic behavior of these immobilized systems is to be analyzed taking into account: (1) The rate equation for native enzymes and for enzymes in gel form is not necessarily the same due to microenvironmental and shear stress effects.
(l-exp(-t/~. Once the layer is formed it behaves like a secondary membrane.. therefore. Therefore. substrate mass balance in this region is expressed according to the dynamics of a stirred tank reactor of volume Vb. all enzyme is virtually confined within region B at a concentration Eb=N/Vb.434
Handbook of Industrial Membrane Technology
in region B can be expressed as follows:
Eh(t)=(k/V. enzyme depth of penetration.” capable of separating compounds of different molecular weight in the mixture as well as catalyzing a chemical reaction. substrate mass balance in region A looks like dSa Sf . Once steady state is attained.Sa . can be assumed identical to the rate equation for enzymes in homogeneous solution. Enzyme immobilization via dynamic formation of an enzyme gel layer has been applied both to flat and tubular membrane reactors.
. For details on the initial conditions and integration of the set of Equations 30 and 31 readers are referred to Reference 30. In most cases.Ra 7a = Ta --dt(30)
The kinetic equation describing region A. that is: dSb Sa . virtually constant over the whole region B. Once steady state conditions are attained. either recirculating the permeate or the axial stream.Sb .)l
(29)
Similarly. xe. In any event. Hence. i. Ultrafiltration of an enzyme solution through a UF membrane does not always result in gel layer formation.)
[I+ . Referring to region B. can then be assumed to be one order of magnitude lower than that for the substrate. enzyme kinetics in region B can be experimentally determined through Equation 32. Unless a gel layer is formed.e.Rb Tb = Tb dt (31) An overall product mass balance at steady state yields the proper rate equation for region B:
Rb = (Q/vb)pb
Pb being product concentration in the permeate. the theoretical model permits calculation of reaction rates in both regions. is equal to the ratio between the corresponding diffusivities. this immobilization technique cannot be used for flow systems lacking effective enzyme immobilization. Substrate concentration is. soluble enzyme membrane reactors can be useful in order to perform kinetic analysis at high enzyme concentrations. it has been shown33 that the ratio between substrate and enzyme depths of penetration.

S) = V. an effectiveness factor. From the fluid dynamics of the reaction vessel. substrate mass transfer and reaction occur simultaneously giving rise to substrate concentration profiles at levels lower than the feed concentration. a mass transfer coefficient. no segregated regions exist: substrate permeation and conversion to products occur throughout the whole gel layer so that both convection and diffusion concur in determining mass transfer through the gel.(Sf .30~32r33~3s The analysis of unstirred membrane systems in flat slab configuration is carried out in two different cases: (1) The gel formed on the flat membrane is homogenous.. an Arrhenius plot) is helpful in assessingwhich step is rate controlling. (3) External mass transfer resistances have to be taken into account occasionally. Thus.Enzyme Membrane Reactors and Membrane Fermentors (2) Within the gel layer.. Since the reaction vessel is unstirred even external mass transfer resistances must be taken into account. (2) The gel is heterogenous. If external resistance is rate controlling. it is possible to estimate substrate concentration at enzyme catalytic sites according to the steady state relationship: K. 77. indicates that kinetics are not rate limiting. Case I: If a homogenous enzyme gel is formed above the membrane surface. A reduction in the activation energy relative to that of the native enzyme.e. KS.“e It is worthwhile noting that most of the available correlations are based upon theoretical models assuming diffusion as the only mass transfer pattern. Hence.e. can be introduced.(Sf-S) = R(S) (33)
which for a Michaelian enzyme reduces to:3 K. namely the Peclet number. the Thiele modulus and the dimensionless Michaelis constant. effects related to external mass transfer resistances are neglected. Expressing equations in dimensionless form pulls out three dimensionless parameters. Comprehensive models of unstirred enzyme gel flat membrane reactors have been proposed. depending on operating conditions and reactor configuration.&Y(K’NI + S)
When the internal mass transfer/reaction step is rate limiting.. is usually introduced related to dimensionless parameters characteristic of the reacting system as a Thiele modulus.
.
435
Experimental kinetic data at different temperatures and at saturating substrate concentrations can be used to evaluate the relative importance of all the aforesaid phenomena.. Plotting such data in terms of the logarithm of the specific reaction rate vs the reciprocal of the absolute temperature (i. substrate steady state mass balance equations upstream and within the gel have to be simultaneously integrated. i. preferential flow patterns exist for the substrate flowing under pressure through the gel.

is negligible relative to cell height. Minimum concentration is.
-=-
dS’
dc’
upstream gel surface dY2 -= dS’ (Dankwaerts conditions). The assumption of a semiinfinite slab for the fluid reacting volume. therefore.1 5’ +m Yl = 1 cell inlet (37)
dY1
dy2
B. In most cases. An overall substrate mass balance relates this factor to the Thiele modulus. xe from Equation 19.C. holds that gel height.C.2 and 3. a:35 ( 1 . and enzymes are uniformly distributed throughout the gel.2
5’
= 1
Yl = Y2
. An effectiveness factor is then introduced. Decreasing substrate concentration is responsible for efficiency factors less than unity across membrane thickness.l.Y2(0) n=(l+M)o2 The existence of convective fluxes rules out the possibility of using correlations already existing. B.C.C. In B. both the gel layer and the supporting membrane are assumed to be completely permeable both to substrate and product. ) (38)
B. defined as the ratio between the overall reaction rate and the rate one would obtain if substrate concentration within the gel were uniform at the feed level. evaluation of the dimensionless parameters is possible and correlations are not necessary.C. the cell output value which can be easily estimated.3’
5’
= 0
0
downstream gel surface
. enzyme kinetic behavior can be assumed to be Michaelian.436
Handbook of Industrial Membrane Technology
d*yl
-+Pe-= dC’*
dyl
0 dE’ upstream (35)
-+Pe-= d(‘*
d*-f2
dY2 d[’ ** Q* /(M + r2)
within the gel
(36)
with the following boundary conditions: B.

3s K E X2 0” = (
) l/2
(41)
D -f*(O) Sf
With reference to the homogenous phase reaction rate. h has to be regarded as the characteristic dimension of the segregated regions corrected by a suitable shape factor: As far as Michaelis-Menten enzymes are concerned. the ultrafiltering surface area per unit reactor volume is quite small. As for membrane UF units. It is interesting to note that transmembrane pressure plays a d. On the one hand. while simultaneous substrate diffusion and reaction occur in segregated regions. increasing pressures lead to increasing permeating fluxes. Within a segregated region. the steady state substrate mass balance equation can be written as:35
-=
d*X d=*
I$.C. Minimum substrate concentration is now attained in the core of the segregated region and cannot be readily measured.1 2
X
(39) (M’+x)
with the following B. thus enhancing reactor productivity.2 z=o dx/dz = 0 center of segregated region (40)
in dimensionless form. high effluent flow rates strongly reduce the substrate conversion. tubular membranes fitted in cylindrical shells in a “tube-and-shell” configuration help in improving the performance of en. Modeling of unstirred UF cells equipped with flat membranes and interpretation of experimental data is then relatively simple. substrate convection occurs through preferential pathways. The modified Thiele modulus.ifferent role from that which it plays in the usual UF membrane separators. @” is expressed as a function of the characteristic dimension of the segregated region and external substrate concentration. They can.29 Moreover.Enzyme Membrane Reactors and Membrane Fermentors
437
Case 2: If gel structure is assumed heterogenous.s: B. and to the feed substrate concentration. therefore. on the other hand. 77 vs @ diagrams have been produced’es for various immobilized enzyme configurations. the effectiveness factor can be estimated to be:3s
dx dz z=l
n =(l+M) Q For preferential pathway geometries other than the flat slab model.C. be used provided that the geometry of the segregated regions is defined and that external resistances are taken into account.C.1 z= 1 x= 1 segregated region/preferential flow pattern interface B.I2
.

immobilized enzymes behave in a manner almost identical to their behavior in homogenous solution. In the second case.40 have been used as cogelling agents. When the enzyme is product inhibited and the effluent from the reactor is recycled.33t4i In both cases. industrial operations require the enzymatic reactor to be continuously operated. High shear stresses may in fact develop on the gel layer surface leading to partial or total removal of the enzyme. Under ultrafiltration at suitable conditions any proteic solution can give rise to a gel layer. In all instances. two other gel formation procedures have been proposed. For example. The enzyme gel layer can also be protected by applying layers of water soluble or insoluble macromolecular compounds or by forming the gel within the porous structure of the membrane where it is less subject to shear stresses. before the UF step.j7 Moreover. under a fluid shear field enzyme molecules can be oriented and thereby denatured. or flushing them along the membrane wall until the gel layer is eventually formed.34f36 An enzymatic gel layer can be built up on the inner wall of tubular membranes either by filtering proteic solutions in a batch mode. The ratio of filtering area to volume is an order of magnitude higher than that for membranes in a flat slab configuration. Under such conditions. In both cases. proteins whose kinetic behavior is affected by the presence of particular compounds in the reaction environment (ligands). This goal can be achieved. Feeding an axial substrate stream to the reactor gives rise to new flow dynamic conditions.2gr32~35~37~39 urease. independent of the nature of the polymer. solutions consisting of both enzymes and high molecular weight inert compounds are ultrafiltered through semipermeable membranes. product accumulates in the feed stream thus inhibiting gelled enzymes. and the flow dy namic conditions are more easily controlled. In the first case. work is in progress to guarantee mechanical stability of the gel layer.37t39 water soluble and insoluble compounds. first of all.31r34r36 A number of procedures have been suggested in order to improve enzyme gel kinetic and mechanical stability. enzymes are chemically linked to high molecular weight inert substances by means of bridge molecules. Besides simple enzyme ultrafiltration. Flushing the substrate solution along the enzymatic gel causes the substrate to be converted to product even in the axial stream. the axial flow rate needs to be optimized since it plays an opposite role.438
Handbook
of Industrial
Membrane Technology
zyme gel reactors.29 p-glucosidase. Cellulosic and polyamide polymers have been used as supporting membrane matrices. High axial flow rates may reduce conversion of substrate to product in the axial stream and enzyme inhibition. namely cogelation and copolymerization/gelation.30p32p33 acid phosphatase. the enzyme microenvironment in the gel layer is characterized by a relevant protein concen-
. 3e In order to develop this immobilization technique for large scale and industrial applications. by strictly controlling the axial flow rate. Yeast invertase. as they do when subjected to less gentle immobilization procedures. at least under laminar flow conditions. 34f36There is evidence of stable systems. polyalbumines (inert proteins). while product conversion in the permeate remains unaltered at a given transmembrane pressure.29 dCMP-amino hydrolase31 and malic enzyme34t36 have been immobilized in gel form on both flat and capillary membranes. Neither allosteric nor pseudo-allosteric enzymes. 2gry1@. show different kinetic behavior.

(A) 16 mg. Deactivated enzyme may be easily replaced.37 tration.Enzyme Membrane Reactors and Membrane Fermentors
439
a 1 0 20 40
Time.
hr
60
Figure 7.25). enzyme gel reactors are inexpensive and easily controlled. Series reactions can be performed in such a set-up.
MEMBRANE
SEGREGATED
ENZYME
REACTORS
In these reactors. Enzyme kinetic stability is enhanced as protein concentration increases (Figure 7. products from one enzymatic layer being fed to the following ones for further transformation. It may happen that enzymes involved in a given transformation cannot be subjected to the same immobilization procedure.25: Effect of gel layer thickness on the stability of cogelled acid phosphatase at 40%. copolymerized/gelled enzymes can lose most of their original activity. The segregation of biocatalyst in the reaction vessel is achieved by means of an ultrafiltration or microfiltration membrane
. provided that it also gives satisfactory mechanical stability. Furthermore. (0) 4 mg. The flexibility of enzyme gel layer reactors is fully exploited when multienzymatic reactions are to be performed. (0) 8 mg. Sequential enzyme gel layers can then be built up on the surface of a membrane in the proper sequence.37 Due to the covalent linking procedure. enzymes or cells are not immobilized. Human serum albumin (HSA) polymers: (0) 2 mg. the deactivation rate of cogelled enzymes appears to be less temperature-dependent than that of copolymerized/ gelled preparations.37r106 enzyme settlement is then particularly favorable. Copolymerized/gelled and cogelled enzymatic layers appear to be mechanically stable over a fairly wide range of temperatures and flow rates in the laminar regime. These features suggest that in many cases immobilization by cogelation is preferred.28t41 In addition. but only confined to a well defined region of space.

over the outer surface of the fibers. Bundles of such loaded fibers were then assembled in a “tube-and-shell” reactor configuration (Figure 7. by capillary action for instance. There is also growing interest in therapeutic applications of compartmentalized cells or microsomes functioning as a bioartificial pancreas or extracorporeal detoxification device. and then sealing both ends.
out
in
Figure 7. Reactors with Enzymes Segregated in the Lumen of Hollow Fibers In 1971. Biocatalytic reactors can segregate enzymes or cells either within the hollow fiber lumen. enzymes and bacterial cells are not lost in the effluent stream.& ” Cylindrical microcapsules” were prepared by filling the core of hollow fibers of a suitable molecular weight cut-off with an enzyme solution.42~45r47~53 or Segregated enzyme reactors avoid the negative aspects of immobilization techniques such as steric hindrance and enzyme deactivation due to coupling. have been extensively used in ultrafiltration units as enzyme or cell reactors for a number of applications.
..46 within the shell surrounding the outer surface of the fibers.26:
“Tube-and-shell” reactor configuration.42 sugar production. Rony first suggested “immobilizing” enzymes within the lumen of hollow fiber membranes. i.43 cellulose hydrolysisW and steroid bioconversion4’ seem to be the most promising industrial applications of such systems. Organic synthetic membranes. Starch and maltose hydrolysis. Enzyme segregated membrane reactors retaining the biocatalyst in the reaction system offer an economically attractive alternative to the design of continuous-flow equipment. Substrate conversion into products can be accomplished if membranes are permeable to both substrate and products. and low-molecular weight products and inhibitors can be removed through the membrane.26) flushing substrate solution into the shell-side.440
Handbook of Industrial Membrane Technology
with a suitable molecular weight cutoff. either in hollow fibers or in flat slab configurations. In this way.s7* The development of hollow fibers with diameters down to about 100 microns makes possible “tube-and-shell” reactors with a high surface-to-volume ratio.44~451575g within the porous membrane support.e.

-
B. so that it is equidistant from six adjacent fibers. a hexagonal free shear layer should sheath each fiber. (4) The shell region where substrate solution flows.1 dv --J (43) = 4k
(r’
r’ = d/a
V = 0 dv -0 ./a
.C.27: Asymmetric hollow fiber schematic: radial cross section. In the absence of pressure driven fluxes through the membrane and assuming a laminar flow pattern. the hexagon is approximated by an equivalent circle of radius Re..
Sponge
Figure 7. which refers to asymmetric hollow fibers. each fiber is assumed to be at the nodal position in an equilateral triangular mesh. (3) The porous spongy part of the membrane.r’ dr’ B. Four different regions can be distinguished in the reacting system (Figure 7.i. 47 In the Waterland model.27) : (I) The core of the fiber where enzyme solution is retained.
(2) The dense thick layer.47
The equations of fluid flow in region 4 are then similar to those of a liquid film falling around a cylinder: 1 d .C.Enzyme Membrane Reactors and Membrane Fermentors
441
A theoretical analysis of such enzyme membrane reactors was carried out by Rony46 and Waterland et al.2
r’ = r.

8’. while the substrate diffusion coefficient in region 2 was assumed one order of magnitude lower. b/a.04x3
P
I#
10-2
10-I
I Thiele Modulus. Accordingly.47 Partition coefficients were kept equal to unity and geometrical parameters were set according to the dimensions of a typical hollow fiber.28: Reactor dimensionless length (z') as a function of the Thiele modulus.
IO _ N
. DJDs. negligibly affect reactor conversion. the reaction rate becomes diffusion controlled where the dimensionless length required for a given conversion is independent of x”. are then a function of seven dimensionless parameters ti2. 3 and 4 were assumed equal to each other.MI
. Essential predictions of the model have been estimated by Waterland et al in terms of substrate bulk concentration as a function of x2. and z’.28) helps in predicting reactor performance at different operating conditions.0) I
I
I
I
I
I
-360
. Operation of such enzyme reactors in a totally or partially diffusion controlled regime can therefore lead to an incorrect determination of biocatalyst stability. if a reactor is operated in a diffusion controlled regime enzyme activity decays. d = 175 pm. and therefore substrate bulk concentration. that is a = 100 pm. z’.I62 .444
Handbook of Industrial Membrane Technology
The set of the above equations leads to a nonlinear problem which can be solved by an iterative numerical technique?’ Substrate concentration profiles.I37 IO1
on0 . The parameter X is substrate conversion4’
.5 pm. down to 10% or even to 1% of initial activity. 0’. d/a.
IO (2 x
I02
lo3
Figure 7. As the Thiele modulus increases. the main dimensionless parameters. A plot in shell coordinates of dimensionless length z’ vs the Thiele modulus x” with 8’ and the degree of conversion at given values (Figure 7.1.391 . DlIy’D2. I 8’. Under these conditions an increase in the activity or the amount of catalyst has little or no effect on the dimensionless length necessary for a given conversion. diffusion coefficients in region 1.237 . b = 100.

4749 Plots of substrate bulk concentration vs the dimensionless reactor length4’ show how steep the descent of profiles can be when large Thiele moduli are approached. and the enzyme solution can be easily recovered and/or recycled. on the initial substrate concentration. Transition from one regime to the other and the Thiele modulus value above which the reaction is diffusion controlled is strongly dependent on the parameter 6’. Substrate conversion is therefore more rapid in the diffusion controlled regime. and the bulk of the feed solution. Furthermore. The production of compounds which alter the pH. in this case.49 An enzyme solution is. Operating the reactor at high feed flow rates leads to a reduction of mass transfer resistances.48 and the partition properties in hollow fiber membranes can result in creating such pH gradients. The term asymmetric refers to membranes comprised of a porous spongy wall supporting a very thin dense layer. Experimentally. Enzyme proteins can be easily retained within the core of the fibers with no deactivation due to coupling agents or to shear stresses. approximately 0. where the reaction takes place. the reaction becomes kinetically controlled.49 Reactors with Enzymes within the Pores of Asymmetric Membranes Asymmetric synthetic hollow fiber membranes designed for use in ultrafiltration/dialysis processes can provide an interesting support for immobilizing enzymes. introduced into one of the membrane-separated chambers of a flat-membrane dialyzer. Therefore. uricase. and reactors are better operated in this regime. The thin skin layer. it can be convenient to operate at small dimensionless lengths at the expense of lower conversions. the technique offers several advantages. more than compensating for this reduction in conversion. As Sf increases. Evaluation of stability and catalytic properties of the immobilized system must take into account possible pH differences between the inner core of the fiber. like ammonia produced from urea via immobilized urease. the region under kinetic control is shifted towards larger Thiele moduli. glucose oxidase and creatinine kinase.2~3~~ Urease.Enzyme Membrane Reactors and Membrane Fermentors
445
As the Thiele modulusdecreases. these differences produce more or less pronounced shifts in the optimum pH dependence of enzyme activity relative to its free form dependence and thereby affect the activity of the enzyme at work. Product mass flow rate will increase. Given the availability of hollow fiber membranes adequately permeable to substrates and products. The overall behavior of such systems appears to be similar to that of the hollow fiber enzyme reactors. in applications where products can be easily separated from the effluent stream.48 alcohol dehydrogenase. uricase. and the control of fluid flow all around the fibers in the bundle in order to assure uniform flow distribution and to avoid stagnation (in order to reduce mass transfer diffusional resistances). Similar reactor configurations using flat membranes in place of hollow fibers have even been used with urease.. except for an apparent higher efficiency.e. increases in catalyst activity or in the amount of catalyst are effective only at high substrate feed concentrations.5
. i.” and alkaline phosphatasesl have been used in such reactors. which can also be achieved using suitable turbulence promoters.

The use of these fibers in bioreactors can eliminate some of the disadvantages of the previously discussed reactor systems.29). the enzymes are effectively immobilized or segregated within the spongy annular section.s3 Enzymes can be entrapped within the outer sponge layer of the fibers by saturating their pores with an enzyme solution (Figure 7. The wet catalytic fibers can then be organized in bundles and assembled in a “tube-and-shell” reactor configuration.29: Schematic of the cross section of a hollow fiber in whose macropores enzymes are immobilized. determines the separation properties of the membrane with little hydraulic resistance to mass transport.446
Handbook of Industrial Membrane Technology
I. if gaseous. This porous spongy wall acts as a mechanical support for the semipermeable thin layer. approximately 75 Pm thick. Diffusion through the membrane matrix and within the flowing solution.42
The dynamics of substrate conversion therefore depend on enzyme kinetics as well as on mass transport conditions. The reactor can then be operated by feeding substrate solution through the lumen of the fibers. play the most important roles in transport mechanisms.
Figure 7. or. The rate-limiting step in sub-
. If substrate molecules are small relative to the membrane’s molecular weight cut-off. Since the flow is laminar in most cases. they diffuse across the dense layer from the fiber lumen into the enzyme solution where the reaction takes place. If the pores in the dense layer are small enough to retain enzyme molecules but large enough to freely pass substrates and products. with pore sizes in the range of IO to 2000 a. and 80 to 90% porosity which has a very high hydraulic permeability.42t4T. product molecules diffuse back into the substrate solution. substrate and product transport resistances through the dense layer are exceedingly small relative to diffusional resistances in the flowing solution. In turn. with pore sizes of 5 to 10 Pm. This skin layer is fully surrounded by a spongy structure.tm thick. leave the fiber through the shell side. in a reverse configuration than that previously sketched.

with substrate solution flowing in region 1. It is assumed we have: 0 0
l
Laminar flow in region 1.)
ar ar
=R
(50)
as2 'D2 . Mathematical modeling of such reactors has been extensively investigated. Vmax
K’M +
R
=
S3 s3
.
l
Using a set of cylindrical coordinates as in Figure 7.. the set of mass transfer differential equations defining substrate concentration in each region is:
i
D3 -r
a
(r as. with substrates and products diffusing through all three regions Steady state conditions Chemical reaction confined to region 3.Enzyme Membrane Reactors and Membrane Fermentors
447
strate conversion is therefore either simple diffusion or the intrinsic kinetics of the reaction itself.(r ) = O r ar ar
1
a
(51)
1
a (r ar asl)_EL ar ax (52)
-~ D1
r
Assuming a laminar flow pattern.) a2
(53)
with enzyme kinetic behavior still modeled by the Michaelis-Menten rate equation. the reaction system may be divided into only three regions. 2. the radial velocity profile in region 1 can be expressed as: r2 v(r) = vmax (1 . 47f53 Referring to the model presented in the previous section (Figure 7. where fluid is assumed to be stagnant Constant fluid properties. and 3.30. namely the previously defined regions 1.271.

The substrate solution is fed to the lumen of the membrane under laminar regime. both along the lumen and the radius of the fiber. Regions under diffusion or kinetic control are approached as the Thiele modulus is increased or decreased. Lewis et aIs and Davis”’ have proposed an analytical solution to the problem in the case of low feed substrate concentration.e. to seven dimensionless parameters: Thiele modulus. When the dimensionless Michaelis constant 8’ (i. 6’ and the following dimensionless quantities: b/a. An iterative numerical solution for the nonlinear problem was fully developed by Waterland et al. Kinetic control holds for larger Thiele moduli. General correlations of bulk concentration as a function of Thieie modulus. X12.. transition between the kinetic and diffusion controlled regimes occurs over a narrower range of the Thiele modulus. d/a. one might say that increases in feed concentration require sensible increases in the reactor dimensionless length to achieve the same conversion as before the change.the reciprocal of the feed concentration) is decreased.47
The intrinsic nonlinearity of this set of equations does not permit an analytical solution.47r53 Plots of the dimensionless length required to achieve a certain given conversion as a function of the Thiele modulus.31).30: Axial section of an asymmetric hollow fiber. that is for a linear rate equation. z’.448
Handbook of Industrial Membrane Technology
-----
x
Ultrothin
Membrane
Figure 7.47 Dimensional analysis of the equations relates the concentration profile. dimensionless Michaelis constant. dimensionless Michaelis constant and the dimensionless length in the entire range of practical operating conditions are available under the same assumptions as the previously described reactor configuration. dimensionless length. respectively (Figure 7. Comparisons between the performance of the reactor operated in this configuration and in the reverse one evidence 47 that both behave in a similar way
. and the slope of thelinear dependence of the dimensionless length z’ on the Thiele modulus (in a log-log plot) increases. From another point of view. Dr/Ds and DI/i’D2. are qualitatively in agreement with those obtained with the reactor operated in the reverse configuration. with the dimensionless Michaelis constant as a parameter. An increase in feed substrate concentration may therefore shift reactor operation from a mainly diffusion controlled regime to a mainly kinetic controlled regime. with a consequent decrease in reactor conversion.

When complex kinetics are involved in substrate conversion. Many enzymes can even be coimmobilized with the macroporous region of asymmetric membranes. as with product-inhibited enzymes (like amyloglucosidase which catalyzes maltose conversion to glucose4*) or with reactions involving a number of intermediates (like starch hydrolysis by means of amyloglucosidase4*) definite information on enzyme kinetics is rarely available. the model applies without modification. when solutions of macromolecular compounds. This reactor configuration is particularly attractive since substrates are physically separated from the enzyme solution only by a very thin membrane layer. 43 Agreement between experimental results and predictions of the more comprehensive proposed model suggeststhat it might be used to predict reactor performances in the case of simple kinetics and when the kinetic and transport parameters are known. The use of small diameter fibers leads to large area-to-volume ratios.g. The second possible effect of the radial flux is to remove enzymes from the fiber wall. that is the ratio of the diffusion coefficient to the wall thickness. Should it happen.. are fed to the reactor.54 invertase.450
Handbook of Industrial Membrane Technology
under conditions of kinetic control. Different operating conditions may require some modification of the performed analysis. A thorough knowledge of the chemical mechanisms through which substrates are converted to products and of the coupling of transport phenomena to enzyme reactions appears to be a prerequisite for the design of such biochemical reactors. e. the dense skin. like starch. diffusional resistances are more pronounced and may hinder the possibility of using such reactors for analytical purposes. Moreover.54 glucose isomerase. Ultrafiltration and/or osmosis can promote convective solute or water flux through the membrane wall. radial convection could compete with diffusion as the main substrate and product transport mechanism. a-galactosidase. Experimental work in which enzymes with relatively simple kinetics were immobilized in the sponge of hollow fibers. thus minimizing mass transfer diffusional resistances. Reactors with enzymes entrapped in the macroporous region of the membrane wall appear to be better whenever they are operated in the diffusion controlled regime. The relative importance of the two transport mechanisms can be evaluated by comparing the radial convective velocity to the diffusive velocity. membranes act as selective barriers protecting enzymes from macromolecular contaminants such as proteolytic enzymes. 47~53 Quick graphical procedures are available in the literature for evaluating the extent to which external and internal diffusion affect immobilized enzyme kinetics. with high enzyme loading capacity per unit of reactor volume.55 and urease.53 are in good agreement with predictions of the theoretical model. A number of other advantages make this reactor configuration feasible for many applications. resulting in the reduction of reactor efficiency. When the Thiele modulus is increased and the enzymes loaded into the sponge of the hollow fibers. When the first one is negligible relative to the latter. the dimensionless reactor length required to achieve a given conversion is smaller than that necessary for a reactor where enzymes are encapsulated within the lumen of the fibers themselves. or selecting substrates on the basis of their permeability or electric charge. The enzyme microenvironment is fully shear free. The procedure itself does not exclude the possibility of crosslinking enzymes directly to the porous polymeric matrix of the fibers. nor that
. Moreover.

The set of differential mass transfer and continuity equations defining substrate and product concentration in these regions are equal to those previously examined. the problem can be overcome allowing the biocatalyst to stay on the shell side. Differences in the environment surrounding each fiber. Apparently the main transport mechanisms through which substrate conversion takes place are: 0
l
Diffusion of substrates from the bulk fluid phase to the membrane Diffusion of substrates within membrane pores Diffusion of substrates within the shell-side of the reactor to the biocatalyst
0
Referring to one fiber alone. the scheme of the reacting system is similar to those examined so far. and the ultrafiltration fluxes make both the analytical and the numerical approach quite difficult. In this reactor configuration. once products are formed they diffuse back towards the stream flowing in the core of the fibers. while the substrate solution is kept flowing within the lumen of the fibers.57 Even though diffusion appears to play an important role in substrate and product transport. 2 or 3. The description of mass transport in the shell-side region is somewhat more complicated. and pure diffusive fluxes. the position of fibers in the bundle. Such binding procedures might in fact result in a greater stability. In some particular configurations. avoiding enzyme purification.Enzyme Membrane Reactors and Membrane Fermentors
451
of compartmentalizing within the same region enzymes previously coupled to soluble polymers or inactive proteins. whole active microorganisms can be segregated in a def-
inite region of space by means of membranes in order to catalyze specific reactions. Microsomess7& and bacteria 61 have been and are currently employed in membrane reactors in order to perform complex multienzymatic reactions or to reduce overall reactor costs. the reaction does not occur in either regions 1. as in the artificial pancreas. Nevertheless. there is experimental evicence that the bundles of hollow fibers assembled in a “tube-and-shell” configuration respond more quickly than could be predicted by assuming purely diffusive fluxes across the membrane walls. that is a slower decay of enzyme activity with time.62 Pressure drop along the length of each fiber should therefore produce a transmembrane pressure across the membrane wall such that at
. When the size of the biocatalyst exceeds the dimensions of the pores in the sponge of the asymmetric membranes.57f63 In cases where a quick transient response is needed.44 previously examined reactor configurations are not applicable. assuming good Models proposeds7~58~62 mixing conditions both in the shell and within the lumen of the fiber.
Tube and Shell Membrane Reactors with the Biocatalyst on the Shell Side
As with enzymes. ‘often deal with simpler systems. especially in the case of membrane units using one large hollow fiber alone. However. such assumptions hold and the models work quite well. the only way to circumvent the slow transient behavior of the device is to reduce the volume where the catalyst is compartmentalized. Such configuration is therefore feasible for all those applications in which products at relatively high concentration are tolerated in the circulating stream.

At a distance. AII. and at the outlet it promotes a backward flux from the shell towards the lumen of the fiber (there is only a small pressure drop in the shel144). the overall fiber-to-shell ultrafiltration net flow rate can then be obtained upon integration of the flux equation over the length of the fiber from the inlet to Lc. can be related to local transmembrane pressure. Assuming that the feed solution is newtonian with a laminar flow pattern. according to the following equation:
J(x) = L
P
Ag (x1
(54)
where the coefficient Lp is the hydraulic permeability of the membrane to the substrate solution.32). 8 = gbf _ IIcnc . the pressure difference promotes an ultrafiltration flux towards the shell side. A quantitative analysis of such ultrafiltration flux59J62 can be approached referring to a single fiber device (Figure 7. say Lc. transmembrane pressure is nil.452
Handbook of Industrial Membrane Technology
the inlet of the reactor. where the catalyst is. pressure in the fiber at distance x from the fiber inlet can be evaluated from Poiseuille’s law to be: 8PQ 1 x (56) np = nIbi . the local transmembrane pressure can be expressed as:
ANx) . Fiber-to-shell solution flux from that point on is negative and becomes a shell-to-fiber flux. that is:
. Neglecting the shell pressure drop.IIs 8lJQ n a4 N f
Introducing
and
C -_
and substituting the expression for flux as a function of transmembrane pressure one obtains:
J(x) = La (6 . x.
(55)
where TIonc and I’& are the local oncotic pressure and shell-side pressure. from the inlet of the reactor local ultrafiltration flux. These pressure profiles promote fluxes which improve the performances of the system as compared to those exhibited by pure diffusive reactors.rib(x)) .flonc -
n.( n a4 N f where flbi c1 Nf = inlet pressure
= substrate solution viscosity = the number of fibers in the bundle. In turn.Cx)
(57)
A decrease in hydrostatic pressure along the fiber due to resistance to substrate solution flow occurs so that at a definite distance from the inlet. J.

Lc)* C and
8*/C = (L . Lp. AII. The use of cofactors with the enzymes in continuous flow systems has of-
. The previous equation can be rewritten to relate the charresponse time of the device and the volume where the biocatalyst is to give: V = A LP All r /8 (64)
Given the membrane surface area.Enzyme Membrane
Reactors and Membrane
Fermentors
453
L
Qfl = 2 II a Nf LP
I
‘(B 0
. and pressure drop along a fiber.. as follows:
L-Lc
Qf2 = -
2 ‘IIaNfTAP I
0
C x
dx = .s9 This reactor configuration is often appropriate for complex catalytic systems.C x)
dx =
(58)
L2 =2aaN f L (BL P
C
-C-‘)
2
But at a distance quently:
Lc from
the inlet ALI = 0. A. upon integration of the flux equation between Lc and the reactor outlet. B* /C = II a Nf LP (L . and pressure drop along the fibers: gfl The shell-side acteristic confined = A LP AIT/ (63)
characteristic time of this reacting system is 7 = V/QfI. the flow rate Q can be evaluated from Poiseuille’s law to give a relation where the net ultrafiltration flow rate. therefore
Qfl + 9f8 = 0 C
(61)
II a Nf L. where V is the void volume.Lc)*
L = 2 B/C = 2 Lc
in the expression of the overall fi-
The expression for B can be substituted ber-to-shell flux to give:
Qfl = 2 P LP L* Q/ a3
(62)
Upon imposition of the pressure drop along the fibers.n a Nf LP (L . is related to the total membrane area. so that Lc = B/C and conse-
Qfl
=naN
B2 L f PC
(59)
The overall shell-to-fiber flux can be obtained in a similar way. QfI. A. it is possible to estimate the minimum volume required in order to obtain the necessary response time. its hydraulic permeability.~c)*
c
(60)
The shell volume or in other terms that is
is constant.

but allow the passageof low-molecular-weight products or inhibitors. This concept has been used in the enzymatic hydrolysis of celIulose. the inhibitor concentration in the reacting stream can be kept low without affecting the overall sugar concentration. Many procedures have been suggested in order to confine these low-molecular-weight compounds in a well defined region of space where they are continuously used and regenerated. respectively. cellobiose (an intermediate product) is formed which acts as an inhibitor for some cellulase enzymes.59
ten been limited by the need to supply large amounts of fresh cofactor. cofactor costs are reduced. thereby keeping their concentration low in the effluent reacting stream. No transmembrane flux occurs at membrane half length.44~65A blend of cellulose. The occurring absorption and reabsorption fluxes along membrane axis are QfI/Qf2. a suitable enzyme mixture. where mammalian liver mi-
. and cellulase.32: Schematic of a hollow fiber encased in a closed shell. In the presence of a suitable regeneration system. only low cofactor additions are needed to maintain excess concentration levels thus assuring maximum rates of conversion. such as the enzyme/3-glucosidase440r the cells of Hansenula. As soon as cellobiose is produced.65 in the shell of the reactor can reduce the extent of inhibition. Bundles of hollow fibers59r66or a single large hollow fiber57r63 in a cylindrical module are used to separate blood flowing within the lumen and mammalian pancreatic islets as assistance to diabetic patients. They have also been sug gesteda for use as extracorporeal blood detoxifiers. substrate. In this way.103r111When the biocatalyst is compartmentalized in this way.454
Handbook of Industrial Membrane Technology
Figure 7. Membranes which retain high-molecular-weight biocatalysts in the shell. is kept flowing within the lumen of the fibers assembled in the bundle.44 Reactors in this configuration are also employed in therapeutic applications. it flows towards the shell being converted to glucose by flglucosidase. usually an expensive compound. When cellulose is converted to glucose.60j64. may also allow the segregation of biocatalysts in the shell capable of transforming inhibitors into nonactive compounds. The presence of a suitable biocatalyst.

as compared to that of the native enzyme. product and substrate mass transfer occur mainly by a diffusive mechanism. Microbial enzymes maintain their activity in conditions otherwise denaturating. Diffusion and ultrafiltration fluxes due to pressure drop along the length of fibers play the most important role in substrate and product mass transfer when systems are operated as previously described. and the fluid dynamics of the solution in the core of the fibers.6g Cell. membranes have been cast with microbial cells incorporated in the casting solution. generally denature enzymatic proteins. It is noteworthy that cell entrapment can enhance microbial activity as compared to cell behavior in homogenous solution. Such immobilization techniques are seldom useful for microbial cells due to their large size. if the membrane molecular cut-off is carefully chosen. presumably because of cellular membrane protection.44 Better reactor performances should result from such operating conditions. When enzymes are entrapped within the sponge of asymmetric membranes. In such systems.61 an effect probably due to cellular membrane permeabilization as a consequence of the entrapment procedure. moreover. once the enzyme is deactivated. Stable immobilized cells permit the carrying out of an enzyme reaction in one stage alone. reactor performance is then controlled only by means of the amount and kind of charged enzyme. an ultrafiltration flux can be promoted from the lumen of the fibers outwards and/or from the shell inwards. The presence of nonaqueous solvents in casting solutions. as well as in polyurethane foams. Enzyme crosslinking in the membrane pores can reduce these losses. Recently. it makes the reactor useless for further operation. and the high temperatures required by membrane annealing. have been discovered which can withstand both high temperatures and organic solvents. but it can lead to an initial activity loss. Sulfolobus solfataricus. result in enzyme losses. However.67 the resistance to flow is fairly low and is strongly dependent on the forming procedure. Entrapped whole cells are the source of a number of microbial enzymes useful for industrial purposes. e. Perspectives Most of the membrane segregated enzyme systems previously examined suffer some constitutive drawbacks which limit their yield and area of application. To date. with a loss in their catalytic activity.. the casting of UF and RO membranes charged with enzymes or whole cells on an industrial scale has been limited by the drastic conditions under which the synthetic membranes are usually formed. the possibility of long term operation make this immobilization procedure extremely attractive. In addition.Enzyme Membrane Reactors and Membrane Fermentors
455
crosomes are compartmentalized in the shell of the module. Of course. membranes will protect transplanted cells from the immunodefensive action of leukocytes in the blood. an op-
.g. loaded membranes appear to be kinetically active and stable over a long period of time. UF or RO fluxes. a number of microorganisms. 61 Polysulfone 61r67and cellulose acetate membranes68 have been cast with microbial cells in the casting solution using the phase inversion technique. In recent years. The main drawback of membrane reactors in this configuration is the relatively slow response to metabolic stimulation due to large shell volumes required to accommodate a number of large-sized biocatalytic units needed to perform the assistance action.

For the sake of simplicity let us assume that asymmetric membranes are used. the mechanisms involved in enzyme immobilization are not well understood in most cases. ligandine and glutathione-S-transferase. A full description of chemical and/or physical procedures required to make enzymes insoluble is beyond the scope of this chapter and is extensively dealt with elsewhere.“4 enzyme immobilization via covalent bonds has been an established immobilization technique. usually carried out by means of extremely active bridge molecules. Enzymes are covalently immobilized primarily onto the surface of the membrane exposed to the feed solution. asparaginase and alcohol dehydrogenase onto nylon tubes.” catalase.
MEMBRANE
BOUND
ENZYMES
IN CONTINUOUS-FLOW
SYSTEMS
In recent years a large number of techniques have been suggested in the Iiterature for immobilizing enzymes on insoluble carriers. in particular. namely enzymatic conversion and membrane separation. uricase. surface bound enzymes are in close proximity to the substrate solution-thus reducing the mass transfer resistance to that associated with the boundary layer. cross-linking and covalent linking to a water-insoluble matrix. they are no longer available to substrate molecules. that suitable active groups are available on the polymeric surface and that the membrane molecular weight cut-off is such that the active layer is enzyme-impermeable. arginase. Asparaginase. resulting in a irreversible loss of activity as compared to the initial activity of the native enzymes. known as the “active side” of the asymmetric membrane. coupling agent molecules. On the other hand. The resultant bonds are generally extremely stable.72j73 and urease. it has been suggested that immobilization results from reaction between the enzyme free amino groups and the glutaraldehyde aldehyde functions with the formation of an intermediate Schiff base. such as CNBr.71 have all been immobilized onto Cuprophan hollow fiber membranes. ‘13 However. dextranase74 and papain” have been attached to cellulosic and polysulfone flat membranes. or if it leads to an enzyme-carrier network inside the polymer matrix. due to the high binding energy of the covalent bonds. can penetrate deep into the active sites of enzyme protein coils where reaction takes place. such as glutaraldehyde. allantoinase and allantoicase. when protease was covalently bound to diazotized polystyrene.112~113 We will be concerned mainly with enzymes bound to synthetic polymeric membranes via covalent binding. In general. Once these sites are involved in linking to the matrix. When glutaraldehyde is used as a coupling agent. even though their activity is often drastically reduced. A stable attachment can result from ionic binding. When the extent of initial denaturation is acceptable in the economics of the process. When enzymes are covalently immobilized in the
. Since 1954.456
Handbook of industrial Membrane Technology
eration which normally requires two different stages. immobilization on polymeric solid supports or glass beads. enzymes bound to membranes can be used in continuous flow reactors. usually quite small. it is not clear whether reaction between enzymes and polymeric membranes via coupling agents simply results in enzyme attachment to the membrane. In this way. Apparently the immobilization procedure affects the immobilized enzyme activity more than the membrane configuration. or hi/multi-functional reagents.

the reacting solution is fed to the core of the fibers. thus minimizing the overall substrate mass transfer resistance. readers are referred to the
As far as membrane
When the substrate has a molecular weight lower than the membrane cutoff.73 In this analysis.C. weight is low as compared
Reactions of substrates whose molecular to membrane molecular weight cut-off. Analytical solutions have been obtained assuming the enzymatic reaction as the controlling step which implies flat substrate concentration profiles all along the tube. electrodes are concerned. the kinetic parameters for native enzymes are no longer applicable for the immobilzedsystem.-3 Comprehensive descriptions of mass transfer and kinetic effects on the performance of such reactors under different operating conditions are not yet available. 1. When such systems are used for analytical purposes. enzymes can be covalently bound either to the “active side” of the membrane or within the sponge-like substructure of the membrane. the diffusional resistance in the bulk phase must be taken into account. Obviously. Assuming laminar flow. Applications of covalently immobilized systems are essentially:
wall must
(1) Membrane electrodes for analytical (2)
(3)
purposes. A theoretical analysis of a tubular reactor with impermeable inner walls coated with enzymes was carried out by Kobayashi and Laidler76r77 and experimentally confirmed by Bunting and Laidler.Enzyme Membrane
Reactors and Membrane
Fermentors
457
sponge of the membranes. the steady state mass transport equation for the substrate is: 2 i as r2 Vmax (1 . the mass transfer through the membrane also be taken into account. when the substrate concentration at the wall can be assumed to
.-) a2 -
as
-
D (Z+-with the following boundary B. Numerical solutions are also available both for diffusion controlled mass transport.
Enzymatic conversion of macromolecules to lower molecular weight species able to permeate the supporting membrane.C. 3
r ar)
ax
= 0
(65)
conditions: x=0 r=a r=O s = Sf D. 1 B. If the enzyme kinetic behavior is not affected by compounds in the solution to be processed. = R(SW)
s = Sf
where Sw is substrate concentration at the membrane wall. 2 B.C. enzymes are preferentially bound onto the “active side” of the supporting membrane.

A group of enzymes from pig liver cytosol. For these therapeutic applications. therefore their removal should hinder further cell development.. the Michaelis constant. Steady state. laminar flow. This model does not apply to a porous wall tubular reactor. Enzymes are in fact purified from mammalian. ar. enzyme immobilization results in either a uniform fixation of enzymes throughout the membrane wall. Hollow fiber membrane reactors with covalently bound enzymes have been proposed as extracorporeal blood detoxifiers or as devices to reduce arginine and asparagine content in the blood of leukemic patients.. Arginine and asparagine content in the blood of leukemic patients has been recently related to neoplastic cells growth. Mass transfer through this solid phase must therefore be taken into account.458
Handbook of Industrial Membrane Technology
be nil. and isothermal conditions are assumed. and a radial convective term must be included.
(66)
. known as glutathione-S-transferase. They describe the dependence of the overall reaction rate on substrate concentration. when the reaction rate is neither kinetically nor diffusion controlled. their removal from the solid matrix would result in a series of immunodefensive reactions which would alter liver operating parameters. In this case. the enzymatic reaction at the catalytic wall enters into the model as a boundary condition.70t79r80 When symmetric membranes are used or when enzymes are fed to the spongy part of asymmetric membranes. Again. Ex vivo and in vivo experiments circulating blood through these devices generally confirm an effective decrease in blood arginine and asparagine content after a relatively short time of extracorporeal circulation treatment.
a
$r’
L as. we must account for solution losses along the tube length. non-human sources. have been immobilized onto the active side of commercial hemodialyzers in order to reduce the toxin content in blood. and the Damkholer number (the ratio of the reaction rate in the absence of diffusional resistance to the diffusion controlled rate). and they are continuously in contact with circulating blood. These enzymes change the polarity of the toxins through conjugation to glutathione. The resulting shift from hydrophobic to hydrophilic facilitates release through the membrane wall. and in the so-called transition region. the irreversible binding of the enzyme to the membrane is extremely important.‘s Asparaginase and arginase79 have been immobilized in commercial Spiraflow hemofilters in order to support anti-leukemia therapies. or in the formation of a carrier-enzyme insoluble network in the sponge of the membrane. A theoretical model neglecting radial convective transport and the dense layer in asymmetric membranes is available in the literature. the enzymes are assumed to be uniformly distributed and the membrane wail curvature is neglected. In addition to applications in industrial processes.81 The reacting solution is still assumed to be fed to the core of the hollow fibers. . enzymes bound to the active side of hollow fibers assembled in a “tube-and-shell” configuration7’ have been and are under study as extracorporeal or “in viva” devices for use in hepatic failure or to assist leukemic patients.rJ2) . Moreover. Differential dimensionless mass balance equations can be written as follows:
a+. Enzyme immobilization via covalent binding generally meets this requirement.1
i (1 .

and the dimensionless reactor length. the transition region becomes narrower as the feed substrate concentration increases. the reaction is diffusion-controlled. Zkin and Zdiff. In order to reduce computer time. a modified Thiele modulus. (r* “)_1 = "max (d-a) . Q. and the dimensionless feed concentration. the reaction is kinetically controlled.+ o----zar' ass -_=O ar" SL = 1
ass
ar"
r’ = 1. w. define conditions to achieve a given conversion under diffusion and kinetic control. at low reactor moduli. The set of non-linear equations is uncoupled by introducing an effectiveness factor. I# -_arl~2 1 sS
Reactors and Membrane
Fermentors
459
r* 1 + B* ss
= 0
O< r”<l boundary conditions:
(67)
in the annular catalyst layer. r” = 0
r" = 1
2' = 0
The concentration history appears to be a function of three dimensionless parameters. Dimensionless reactor lengths. 0”. now Z’ = DL x /2 vmax a’. y*.
. Results resemble those obtained by Waterland et al in their model of hollow fiber reactors with the enzyme compartmentalized in the sponge of the polymer matrixP’ At a low dimensionless reactor length. latter being equivalent to the inverse Graetz number and proportional to the reactor residence time (see Figure 7. the effectiveness factor has been conveniently expressed as a weighted sum of its value for the zero and first order reaction rate. 2 K'M a
analogous to the Damkholer number. r" = 0
a. the mass Biot number. and a high reactor modulus. the (Y”WP. Different regime conditions are depicted in terms of a dimensionless parameter. with the following
a& -_=O
ar' sL = ss
r' = 0
r' = 1.Enzyme Membrane in the bulk liquid phase. The progress of the reaction is described in terms of the reactor modulus.33). and numerically solved. As in the model proposed by Waterland et al. This is called the reactor modulus and is obtained by collecting the mass Biot number and the Thiele modulus together.

81
.33: Reactor modulus as a function of reactor dimensionless length at various values of the reciprocal of the dimensionless Michaelis constant.460
Handbook of Industrial Membrane Technology
-2L
4
I
I
I
1
I
I
I
I
I
3 2 r3 AI 8 i-0 -I -2
-I
-
p&N
I -2 I -I 0
I
-2 -3
log 2’
Figure 7. The parameter X is substrate conversion.

at a given conversion.81
.34) show how the functional dependence of f$.
-2
0
2
4 -log
6 +I
Figure 7. Horvath et aIS1 also give suggestions for the estimate of a reactor effectiveness factor using dimensionless reactor lengths under limiting kinetic and diffusion controlled regimes. Eoiff is the ratio of the actual rate to the rate that would produce zero substrate concentration at the catalytic surface. Reactor efficiency can thus be related to a kinetic efficiency. the dimensionless reactor length at limiting conditions can therefore be used to define an overall reactor effectiveness factor as the ratio of a minimum length characteristic for the reactor to the actual. and a diffusive efficiency. Ekin compares the actual heterogenous catalytic rate to that achievable with the same amount of enzyme in a homogenous plug flow reactor under the same conditions.. The average reaction rate is inversely proportional to the dimensionless length. on the reactor modulus is similar to that of the catalyst efficiency on the Thiele modulus. This expression simplifies the treatment of data and reactor design considerably. The parameter is the reciprocal of the dimensionless Michaelis constant. The reactor effectiveness factor can be expressed as the ratio of the actual reaction rate to a maximum rate characteristic of the reacting system.34: Kinetic efficiency as a function of reactor modulus at two values of substrate conversion. Ekie = Zkie/Z’. Ediff = &jiff/Z’.. Plots of Ekin vs the reactor modulus (see Figure 7. it is therefore a measure of the degree of utilization of the enzyme.Enzyme Membrane Reactors and Membrane Fermentors
461
Generally the dimensionless reactor length required to achieve a given conversion can be expressed by the approximate relation:
where <v> is an average catalyst effectiveness factor.

Design equations can then be obtained by the definition of limiting reactor length under bulk diffusion control. at low reactor moduli (kinetic control). due to axial changes in the boundary layer resistance. the usual Eadie or Hofstee plots should not be used. this model is no longer accurate. Mazzola et al’l showed that a value of the axial flow rate exists at which the maximum reaction rate is attained. a kinetic analysis has to be performed in order to characterize enzyme behavior after the immobilization procedure. straight lines are obtained and these plots are reliable for evaluating enzyme kinetic parameters. After the enzyme is immobilized. obtained when the reaction is rate-limited by substrate transport in the bulk liquid phase.462
Handbook of Industrial Membrane Technology
When such reactors are used for analytical purposes. The reactor dimensionless length required to achieve a given conversion is then given by the expression for Z under kinetic control: *
zkin
f = -
0
a
(/!I*x . Figure 7. uricase. in any case. The highest reaction rates are. In the diffusion controlled regime. commercial Cuprophan hollow fiber membranes7’ are often used.71R74 In extracorporeal devices. In vivo fluid dynamic conditions strongly limit the range of conditions at which such reactors can be operated. allantoicase and allantoinase have been immobilized by means of glutaraldehyde or CNBr onto these membranes. Reactors are often forced to operate in the diffusion-controlled regime. higher flow rates result in a decrease in apparent enzyme activity. Since kinetic parameters cannot be assumed equal to those of native enzymes.ln(l-X))
When radial convective fluxes due to transmembrane pressure occur. no initial rate can be determined in the conventional sense even when the catalyst effectiveness factor is unity.81 However. catalase. at the same time. Enzyme immobilization in the sponge of polymeric asymmetric and symmetric membranes has the advantage of a stable immobilized enzyme system along with the improved isolation of enzymatic proteins from immunodefensive system actions. high molecular weight inhibitors and proteolytic enzymes. separating undesirable compounds from the enzymatic proteins. A suitable choice of membranes based on their separation properties allows substrate molecules to permeate through the membrane while. it is wise to check the integrity of the membrane (e.35 shows how immobilization of urease by diazotization on heterogeneous polysulfone flat membranes reduces the membrane hydraulic permeao bility by 50/o.g. the overall reaction rate depends on the substrate supply rate. Under these operating conditions. Sometimes even membrane transport and mechanical properties are affected by the harsh chemical treatments required in the immobilization procedure. the permeability to the species of interest and the strength). relatively inactive enzymes are needed. When maximum utilization of the available catalyst is the main goal. The mass transfer mechanisms operative in substrate conversion are essentially those described by Waterland et al47 in their model of the compartmentalized enzyme membrane reactor. 74 It also shows how the higher permeability membranes are more affected by the immobilization procedure. In order to reduce the side effects of free enzymes in intravenous injection therapies. the apparent enzyme activity can thus increase by increasing the recirculation flow rate.. asparaginase. In in
.

35: Volume flow as a function of pressure using two different types of membranes. This evidence suggests that more than one metabolite has to be removed from blood at the same time to make the extracorporeal treatment effective. pure water (0).36). when uric acid is fed to the multi-enzyme reactor.. due to high pressure drops arising from blood viscosity.83 Even though the use of such reactors as extracorporeal units can result in an effective decrease of asparagine or arginine content in blood. 3 to 4 hours (Figure 7. dextran T4 (4000-6000 daltons) 1 wt% (a).‘r While asparagine depletion in plasma is accomplished quickly. Filled symbols: membrane type 3.Enzyme Membrane
Reactors and Membrane
Fermentors
463
0
1
2 3 I Pressure (bar)
5
Figure 7. asparagine again appears in the plasma. and when it levels off at its initial value. Up to four enzymes involved in the metabolic pathways of purine basesallantoinase. This behavior has been confirmed by experiments on leukemic patients. using a 0. Test conditions: uncoupled state.015 M urea solution (o). after the treatment. Reactor performances are depicted in Figure 7.74 vivo experiments. allantoicase. type 2 being more permeable than type 3. its effect is limited to a few hours after the treatment. uncoupled state. i. in blood cells it proceeds more slowly. coupled with urease.e. In vivo extracorporeal circulation experiments on rats with asparaginase bound to the outer surface of Cuprophan hollow membranes assembled in a “tube-and-shell” reactor configuration confirm that even this kind of reactor can eliminate asparagine after a short period of time. long reactors may not attain maximum reaction rates. the asparagine content in blood cells suddenly rises.37. uricase and catalase-have been immobilized together by means of glutaraldehyde on the outer surface of cellulosic hollow fibers. Higher apparent reaction rates result when operating with shorter reactors at higher flow rates but under the same pressure drop. Blank symbols: membrane type 2.‘r
. Unfortunately.

37: Performance of a multi-enzyme membrane reactor when uric acid is fed as a substrate.71
. uricase and catalase arecovalently bound to the outer surface of a cellulosic membrane.
IOC
l
Plasmatic nmoles/ml
L-Asn
50
l
Piasmatic nmoleslml
L-Glu 400
0 Blood cells L-Glu nmoles/ml 200
Figure 7. allantoinase. such as albumin. or polyamines in order to shift the maximum enzyme activity towards blood pH. time in in vivo experiments performed on rats.71
A Uric
acid acid
l
Allantoin acid
CJAllantoic
V Glyoxylic
Q-
I I
2 3 4
I
5
8
6
I
7
P ‘t I
a
2C Time h
Figure 7.464
Handbook of Industrial Membrane Technology
Enzyme stability and activity can be enhanced by performing enzyme immobilization in the presence of inert proteins.36: L-Asparagine concentration vs. The extracorporeal device is a “tube-and-shell” membrane reactor where asparaginase is bound to the outer surface of Cuprophan membranes. Allantoicase.

they are converted to low-molecular-weight products which permeate the membrane. Figure 7.38: Sum curves of molecular weight distribution for different pressures.3874 shows how transmembrane pressure can affect molecular weight distribution in the permeate for the system dextran-dextranase when enzymes are bound to polysulfone flat membranes by means of glutaraldehyde. especially proteins.0. is strongly hindered by the so-called concentration polarization and fouling phenomena (see Chapter 3 on ultrafiltration). Immobilization of proteolytic enzymes onto the active side of the ultrafiltration membrane can provide a self-cleaning function. As soon as macromolecules come in contact with the active enzymatic layer.
100
80
20 0
2
3
4 In MW
5
6
Figure 7. T = 40°C. Ultrafiltration of macromolecular solutions. The hydrolytic cleavage of large proteins by proteolytic enzyme is controlled by the amount of immobilized enzyme. The irreversible deposition ” of a macromolecular gel layer on an active membrane surface usually results in a sharp decay in permeate flux.74
. pH = 6. (17)4 bar.Enzyme Membrane Reactors and Membrane Fermentors Self-Cleaning Enzymatic Ultrafiltration Membranes
465
Thus far we have examined some of the possible applications of enzyme membrane reactors. the fluid dynamic conditions under which the system is operated and by membrane rejection properties. immobilized enzymes can also be used to improve the performance and effectiveness of traditional ultrafiltration processes. focusing attention on situations in which enzymatic conversion is essential and the membranes act only as porous supports. dextran T70 1 wt%: (0) 0. However. (A) 3 bar. (0) 1 bar.5 bar. Polysulfone membrane type 1.

it is possible to prepare semipermeable membranes whose ultrafiltration yields are higher than those of passive membranes. (A) 5 bar.19 wt%.0. saccharose.74
MEMBRANE
FERMENTORS
Increasing interest is developing in continuous fermentation processeswhere microporous membranes are used to separate the fermentation broth from the product stream. dextran T70.”
2
4
In MW
6
Figure 7. dextran T70 1. (0) 3 bar. Blank symbols permeate.39)74 Even though enzymatic conversion is not too effective. thus retaining viable cells in the fermentor. %) in permeate and reactor after 25 hr. By carefully choosing operating conditions. traditional fermentors could in fact be coupled to a membrane separation unit in a configuration called a “membrane fermentar” where flux through the membrane was pressure driven.39: Sum curve of molecular weight distribution (C. macromolecular concentration in the permeate increases (see Chapter 3). (0) standardization solution. it is possible to gain a net shift in molecular weight distribution curves (Figure 7. Polysulfone membrane type 2.466
Handbook of industrial Membrane Technology
At low transmembrane pressure.74r75 Ultrafiltration experiments of cheese whey through cellulosic membranes to which papain was covalently bound. pH = 6. In 1970. As the pressure increases. show that flux decay curves of enzymatic membranes are even lesssensitive to pH changes. Since then a number
. (0) standardization solution. Porter and Michael? first suggested the use of membranes to enhance fermentor productivity. macromolecules are nearly 100% converted to products. filled symbols reactor.

A typical apparatus is shown in Figure 7.40 in which there is a vessel for the growing biomass.% confined within the shell of a hollow fiber memor in a cell circuit separated with a dialysis memwhere water or substrate solution is kept flowings6 models for the growth and fermentation of a speon the corresponding substrate medium are avail-
Viable cells can also be brane module8’ (Figure 7. (Figure 7.85
.Enzyme Membrane Reactors and Membrane Fermentors
467
of different membrane configurations have been proposed.41:
Schematic of a hollow fiber fermentor (HFF). with membranes in hollow fiber or flat slab configuration.42). When kinetic cific kind of cell or microbe
Product
Flowmeter
Hollow
fibre module
Figure 7. to withdraw products from the flowing cell slurry. and a UF unit.40:
Cell recycle fermentation apparatus. where the pH and temperature are strictly controlled and nutrients are added.
LWd
Control _Y
Proba
Ullratiltrolm7
Umt
F
Figure 7.41) brane from a dialysate circuit.

.40). Due to complete mixing. as is the case for Zymononas mobilis. Vessels are usually designed to insure a uniform concentration of nutrients and pH throughout the whole volume. In most cases.42: Schematic of dialysate-feed.2 and 7. 7. The sludge is flushed back to the reactor.1. In Tables 7. a summary of the performance of different fermentor configurations is reported for three fermentation processes. immobilized continuous fermentation. and a membrane separation unit (as in Figure 7.468
Handbook of Industrial Membrane Technology
FERMENTOR CIRCUIT
DIALYSATE CIRCUIT
Figure 7.%
cell system for dialysis
able. carried out by different microorganisms. the reactor biomass is pumped to the UF unit where solid-liquid separation occurs. the flow rate of nutrient feed is kept equal to the permeate flow rate thus keeping a constant liquid level in the anaerobic reactor.8’ mathematical modeling of membrane fermentors can be accomplished according to mass balances and the equations reported for the corresponding enzyme reactor configurations examined in the previous sections.s8 After anaerobic stabilization. process control and stability of the microbial slurry are not difficult to achieve. when the biomass is well developed.3. Ceil Recycle Fermentors Cell recycle fermentors consist of two main units: a vessel where the biomass is allowed to grow.

over the vessel alone is as follows (see Figure 7. and X’ = Y D’ (Sf . For a suspended growth system.88 The specific growth rate (SGR) is equal to the reciprocal of the SRT.
X' = D' Y (Sf .Co’).40):
a’
F Xi .S)/ P*
Combining the two equations.43) can be written as: 11’ F X’ and
=
Ll*x’
v
(68) respectively. Steady state mass balances on viable cells and on the limiting growth substrates for a continuous fermentor (Figure 7.S)
A steady state mass balance on a continuous cell recycle fermentor. making process control easier. When high biomass concentrations are achieved. longer SRTs can be maintained at lower reactor volumes.S)/ p* on limiting growth substrate to give a cell concentration value in the vessel equal to:“’ Xi = Y (Sf S)/(l + a’ C a’) (71)
The cell concentration within the vessel of a cell recycle fermentor is then greater by a factor of l/(1 t (Y’ . l/time. KS.
. In turn.
The kinetics of substrate removal in the anaerobic reactor actually determines the SRT required for a given fermentation efficiency. the presence of the UF unit allows a fine control of SRT. and the reactor volume loading (VL). mass/mass time = the organism decay coefficient. SRT is usually defined as the ratio of volatile suspended solids (VSS) in the reactor to the VSS lost in the effluent or intentionally wasted per day. The SRT is the average retention time of organisms in the reaction vessel.(1 + a') F Xi + V P* Xi
= 0
(70)
on cells. to give:88 SGR = Ye KS-D where Y KS D = the organism yield coefficient: the massof cells formed/mass nutrients = the specific substrate utilization rate. The performance and capacity of an anaerobic reactor can be expressed in terms of two parameters: the solid retention time (SRT). it can be expressed as a-linear function of the specific substrate utilization rate.470
Handbook of Industrial Membrane Technology
Higher cell concentrations are usually achieved in cell recycle fermentors than in the usual fermentor. the cell concentration within the vessel is:ll’
x’ = Y(Sf .

steadily decrease and increase. as in the case of wastewater treatment in the dairy industry. Product and substrate concentrations in the permeate.‘.
’ . Reactor design can be carried out by estimating the value of volumetric loading necessary to achieve a given effluent quality. .e.43:
Schematic of a single stage chemostat. as dilution rate increases. and of productivity on permeate flow rate.* As dilution rate increases fermentor productivity increases. On the other hand. instead. dilution rate.44 shows the typical dependence at steady state of substrate and product concentrations.Enzyme Membrane Reactors and Membrane Fermentors
471
f
(literdr)
:. VL. so the VL or the SRT approach can also be used as design criterion..
s.
*.I
I
L
X (&liter)
S (&liter)
-
4
V
liquid
volume (liters)
_ . attains a maximum value and then decreases. This is not the case with cell recycle fermentors. A compromise generally has to be made between production rate and product concentration in the effluent. When the absence of substrate in the permeate is required. it obviously limits fermentor productivity. This design criterion is commonly used for biological systems where the evaluation of biomass concentration within the reactor is difficult. low substrate concentrations in the permeate keep recovery
. is defined as:59 VL = 0 SflV = Sf/T
High biomass concentrations permit operation at higher volumetric loading rates.“’
The volumetric loading to a reactor.
Figure 7. i. respectively. for lactose fermentation to ethanol by Kluyveromices fragilis in a cell recycle membrane fermentor. Figure 7.
c 3.

to give J = KS log (C. Therefore.44: Fermentation kinetics of a membrane recycle fermentor. Feed concentration = 150 g/Q lactose. The design of the UF membrane unit in the polarized regime must relate the flux to the optimal level of reactor VSS or total suspended solids (TSS) according to Michaels’ gel theory. and increases in the permeation rate can be achieved only by enlarging the membrane surface area or by improved fluid management. These data outline a dominant feature of the cell recycle membrane fermentor: productivity is a monotonic function of viable cell concentration within the fermentor. Data in Table 7. Cell recycle makes it possible to operate at higher microbial cell concentrations than in conventional batch fermentors. Cell concentration = 90 g/P. permeate fluxes become invariant with the transmembrane pressure.a6 When concentration polarization occurs.gOInvariably. thus reducing the volume required for a given productivity and hence the capital costs. the viscosity of the cell slurry increases with cell concentration./%
.” In high rate fermentors. Complex systems such as rotating membrane fermentors have been proposed to overcome fouling and expected concentration polarization problems. so that concentration polarization phenomena are usually significant.se
costs of products from the effluent stream low.472
Handbook of Industrial Membrane Technology
L
-
en
240
‘+
\
ci E:
580 I
ilJ
>
5
0
4 Dilution
8 Rate 6. hr-l
Figure 7. the optimum operating conditions will be determined by the economics of the overall process.1 show how the substrate concentration in the feed can affect fermentor performance. concentration polarization is usually controlled by operating the reaction system at high recirculation rates though pumps have to be carefully chosen to avoid cell damage.

Since the cell slurry is separated from the substrate solution by the membrane. In this way.“’
.g3 Membrane Segregated Fermentors Continuous fermentation processes can also be carried out in a hollow fiber fermentor (HFF). Initial cell concentration = 117 g/L?. primary inhibitors are released through the membrane.s’ Cells are packed into the shell side of a hollow fiber module. HFFs can be used for fermentation of liquid streams containing low molecular weight fermentable substances. (A) ethanol. In ethanol fermentation processes.” The use of a membrane able to remove ethanol selectively’* has been proposed.
473
High recirculation and dilution rates help in maintaining low levels of inhibitory products. Both primary and secondary metabolites freely pass through the membrane thus relieving both primary and secondary inhibition. A careful choice of membrane molecular weight cut-off can also assure that the cell environment is fully sterilized.45: Fermentation kinetics of a hollow fiber fermentor in single pass mode processing whey permeate (45 g/Q lactose concentration). but secondary products are allowed to accumulate. Figure 7.45 shows typical performance of an HFF for whey permeate fermentation to ethanol by
0 2
I
*
1
I 4
I1
I
I 6
I. and = anaerobic reactor VSS or TSS. their concentration level soon becomes toxic reducing cell viability and alcohol productivity. = apparent solid concentration in the gel. (0) productivity.
Figure 7.
* 8 d(hf’)
15
Dilution
Rate.(0) lactose.Enzyme Membrane Reactors and Membrane Fermentors where Ks cg x = the overall mass transfer coefficient. while substrate solution is fed to the core of the fibers (Figure 7.41). for instance. the removal of these inhibitory products through the UF membrane can lead to significant improvements in reactor performance.

reactor productivity is increased at the expense of low substrate conversions.3 Q/h after 80 hours of operation. performances are not yet comparable to those of other membrane fermentors. the cell slurry is separated from substrate solution by a dialysis membrane86 (Figure 7.46) is also better than in a batch fermentor.2). In the second. There are some problems related to attaining steady state in terms of cell growth. Moreover. even with HFF. where it occurs.
0
80
160
240
Time. The long term cell stability in an HFF (Figure 7. However. Even though improvements can be obtained by operating the fermentor according to the first suggested procedure. (a) productivity.”
. (A) ethanol. Substrate entering the fermentor is sterilized by passage through the membrane and its consumption for cell growth is reduced to metabolism maintenance. substrate is fed into a continuous fermentor circuit that is dialyzed against a continuous dialysate circuit where water is kept flowing.7 to 1. substrate is fed into the dialysate circuit and diffuses towards the batch fermentor circuit through a dialysis membrane. Initial cell concentration is 100 g/P.42). Dialysis membrane fermentors provide a novel approach to the immobilization of microbial cells and the relief of both primary and secondary inhibition.1.2). In another membrane fermentor configuration. (0) lactose. 7. Dilution rate (D’) is changed from 2. Two different operational modes for this configuration have been proposed. cells are effectively immobilized. The behavior of the fermentor is similar to that of a cell recycle reactor with the same fermentation system.474
Handbook of industrial Membrane Technology
Kluyveromices
fragiis.% In addition. at least at low dilution rates. In the first.hr
Figure 7.46: Long term stability of a hollow fiber fermenter in single pass mode processing whey permeate. Higher productivity and yields than with a batch fermentation are obtained (Tables 7. the solute exchange capacity of dialysis membranes limits both productivity and conversion (Table 7.

Hong. 9. 23: 1501-1516 (1981). and Majewska. Michael D. Immobilized Enzymes. A. and Bunting. K. J. 6. 7. 4. Wandrey. Immobilized Enzymes: Halsted Press (1978). A comparison of continuous SSF to batch SSF shows how performance can be improved with the use of membranes. 8. K. J. Enzyme consumption is strongly reduced when the reactor is operated in this way. 5.. Immobilized Enzymes. Bioeng.476
Handbook of Industrial Membrane Technology
Since the optimum temperature for the hydrolysis step is usually different than that for fermentation. G. Zanichelli (1976). Biochimica. Proceedings of the International Conference on the Commercial Applications and Implications of Biotechnology.. mobilis cells. Qarendon Press (1973). 577-588.C... Laidler. Mosbach. The continuous SSF membrane process also offers appreciable increases in productivity. Enzymes exhibit their highest activity at temperatures appreciably higher than those required for microbial fermentations. P. a “one-step transformation” may not be as efficient as two steps. two fermentors are connected in series. 3.98 In this process. Typical steady state performance of a continuous SSF membrane process is reported in Figure 7.9’ Interesting applications of these systems can be found in ethanol production from cellulose or starch. The overall process is called simultaneous saccharification and fermentation (SSF). fermentation time and capital costs can be reduced. UF membranes can in fact be used to recycle both enzymes and microbial cells thus increasing overall system productivity.47. and Catapano.
REFERENCES 1.L.S. T.
. J. Biotechnoi. New York. G. Drioli. Tsao.T. and Wankat.. Biotechnol. where simultaneous saccharification and fermentation occur due to the presence of amyloglucosidase and 2. I. The first one provides for the liquefaction of the slurry. Winnicki. When such differences are
negligible.. (though at the expense of reduced product concentration in the effluent stream) as well as decreased energy requirements when compared to a conventional two-step process. Wiley and Sons (1981). 2. Oxford. which is then pumped into the second fermentor.J. E. cwzymatjr hydroJy+is and fermeniatjjon
can shdtaneolldy
OCCIJJ. Chimica Oggi 7/B: 11-16 (1984).. pp. Online Publications Ltd..... P. Lehninger. UK (1983). 25: 1441-1452 (1983). TIBS January l-3 (1980). Bologna. Bioeng. Chibata. Wisniewski. What has been previously said about the advantages of continuous membrane fermentors also applies to such complex systems. Trevan. C. K. Enzyme Membrane Reactor Systems. Northwood. The Chemical Kinetics of Enzyme Action.
Provided that enzyme concentration is high enough to prevent enzymatic conversion from being rate limiting..

The electrical charges of ions allow them to be driven through solutions and water-swollen membranes when a voltage is applied across these media.la with no membranes present.ions are generated at the cathode. but the ions do not have to travel very far to reach the membranes in the thin solution compartments of ED stacks. usually less than 1 cm/min. Davis
INTRODUCTION Electrodialysis (ED) is an electromembrane process in which ions are transported through ion-exchange membranes from one solution to another under the influence of an electrical potential. Ionic transport is illustrated in Figure 8. When NaCl is dissolved in water. Clz gas is generated at the anode while Hz gas and OH. but the products would be impure. it dissociates into Na+ cations and Clanions. Hz and NaOH from the electrolytic cell in Figure 8. reverse osmosis. design emphasis is directed toward maintaining uniform distribution of solution flow and minimizing electrical resistance and current leakage. Electrode reactions. transfer the current from the solution to the anode and cathode. which will be discussed in detail. In conventional electrolytic cells.la. Some important concepts of electromembrane processes are presented in Figure 8. The combined motion of Na’ to the right and Cl. and filtration. The nature of the membranes and the driving force distinguish ED from the pressure driven membrane processes such as gas permeation. Application of an electric potential to the electrodes causes the ions to move through the solution at velocities proportional to the strength of the electric field. the requirements for strength and support of the membranes and containment vessels are less demanding in electromembrane processes than in pressure driven membrane processes.8 Electrodialysis
Thomas A. Instead.to the left carries the total electric current flow through the bulk solution (unless other ions are present).1. Since pressures are usually balanced across the membranes. Average ionic velocities in the bulk solution are surprisingly slow. a porous diaphragm between the electrodes prevents bulk mixing of the products 482
. One could produce C12.

and OH. When a pair of anion and cation-exchange membranes is placed between the electrodes as shown in Figure 8. The degree of crosslinking and the fixed-charge density affect the membrane’s properties in opposite ways. This exclusion. Since their concentration is very low. as a result of electrostatic repulsion.g. Divinylbenzene is used to crosslink styrene. is called Donnan exclusion.ld. crosslinking is needed to prevent dissolution of ion-exchange membranes. Ions with a charge opposite to the fixed charge (counter ions) are freely exchanged at these sites. because the fixed -NRs’ groups repel positive ions.Eiectrodialysis
483
of the electrode reactions. polystyrene). This innovation has substantially improved the purity of the caustic produced. therefore.lb. The anion-exchange membrane allows Cl.g. It is this selectivity for electrolytes that often makes ED the process of choice for certain separations. Na’) is relatively high. The enriched and depleted solutions are withdrawn from their respective compartments to achieve useful changes in the electrolyte content of solutions without substantially affecting the content of nonelectrolytes.g. but it does not allow Na’ ions to enter. This membrane electrolysis cell is expected to dominate chloralkali plants in the future. -SO. Modern chloralkali cells utilize a cation-exchange membrane between the electrodes as illustrated in Figure 8. Higher crosslinking improves selectivity and membrane stability by reducing swelling.1~ were switched..and cation-exchange membranes as shown in Figure 8. desalting of protein solutions or whey..g. Ion-Exchange Membranes The ion-exchange membranes used in electrodialysis are essentially sheets of ion-exchange resins. nearly all of the current is utilized to produce CIZ and NaOH. the solution between the membranes becomes depleted of NaCI. This effect is accomplished in an ED stack by an array of alternating anion. but it promotes swelling and thus necessitates higher
.1~. This depletion occurs because the membranes are selectively permeable to ions of a specific charge. anions carry only a small fraction of the electric current through a cation-exchange membrane. High charge density reduces resistance and increases selectivity. in this case the anions.ions.ions to be carried out of the center compartment by the electric potential. The solutions between the membranes are alternately enriched in or depleted of NaCl when the electrodes are energized. but it increases electrical resistance. Similarly the cation-exchange membrane offers a low resistance to Na’ ions and a high resistance to Cl.g. Therefore. the NaCl content of the center compartment would increase.ions. -NRs+) to the polymer chains forms anion-exchange membranes.2 illustrates the structure of a cation-exchange membrane which has negatively charged groups (e. Attachment of positive fixed charges (e.) chemically attached to the polymer chains (e. e. The cationexchange membrane is permeable to Na’ ion but relatively impermeable to Cl. Figure 8.. The concentration of counter ions (e. Since ion-exchange polymers (e. counter ions carry most of the electric current through the membrane. If the polarity (charge) of the electrodes were reversed or if the positions of the two membranes in Figure 8.. The fixed charges attached to the polymer chains repel ions of the same charge (co-ions).. styrenesulfonic acid) are water soluble. which are selectively permeable to negative ions..g.

Water Transport The permeability of ion exchange membranes to water is seldom tabulated as a membrane property. A long-term study by Kneifel and Hattenbach’ revealed deterioration in some of these properties after prolonged exposure to 0. NaOH tended to be most destructive. The compartments usually contain spacers that keep the membranes separated by a constant distance.1 are for new membranes.’ The membrane properties reported in Table 8. Commercial ED stacks typically contain more than 100 cell pairs between a single pair of electrodes.
ED STACKS Design Considerations An ED stack consists of alternating anion. It also contains connections for solution streams into and out of the enriching and depleting compartments.3 constitutes one cell pair. because it varies so much with types and concentrations of ions. The solution compartments are bounded by perimeter gaskets that are pressed tightly between the membranes to confine the solutions within the compartments. NaOH. An exploded view showing the components of an electrodialysis stack is presented in Figure 8. The two membranes and two gasket-spacers illustrated in Figure 8. water transport determines the upper concentration limit of the enriching stream when ED is used to concentrate solutions. Holes in the inside face of the end plate are aligned with holes in the membranes and gasket-spacers that form manifolds for solution flow through the ED stack. However.3 When solutions are dilute and membranes are loosely crosslinked. because the counter-ion transport number drops as the concentration of the external solution increases. (Poor flow distribution within and among the compartments severely limits the performance of an ED stack. However. the conditions of measurement must be specified. Ions that are transported through the membranes by an electric potential drag along about five water molecules on the average.) Each solution compartment has a means for introduction and removal of solutions. It contains an electrode and the connections for the solution that rinses the electrode. The spacers also aid in distributing the solution velocity evenly throughout each compartment. Most ED stacks are designed to use identical gasket-spacers for enriching and depleting compartments.and cation-exchange membranes with solution compartments between them. Water transport is not often a significant factor in ED of dilute solutions.
. substantially more water transport occurs.3.Electrodialysis
487
The transport number of the counter ion provides an indication of the permselectivity of the membrane.1 N solutions of NaCI. Large stacks have multiple manifold holes on both ends of the gasket-spacers to provide uniform flow over the width of the compartment. All of these features are combined into a plastic device called a gasket-spacer. Techniques for measuring the transport number are described by Helfferich. Membranes that are sufficiently crosslinked to restrict water transport in dilute solutions are referred to as “tight” membranes. and HN03. The end plate is typically made of rigid plastic.

The thinness is ultimately limited by hydraulic resistance and by problems associated with getting solutions into and out of the solution compartments between the membranes.4. To maintain constant spacing between the membranes it is necessary to support the membranes at close intervals.488
Handbook of Industrial Membrane Technology
-
1
ANIONEXC”*NGE MEMBRANE w. Each concentrating cell consists of one cation-exchange membrane and one anion-exchange membrane sealed at the edges to form an envelope with a spacer screen inside. The lonics design utilizes the tortuous-path spacer illustrated in Figure 8.N”
MORE
MEUBRANES. These supports are more widely spaced than those of a sheet-flow design. Inc. Superimposed on this channel is a spacer that supports the membrane and promotes turbulence in the solution compartment. Unit-cell stacks were specifically developed for concentrating solutions.MES. is the major exception). but the support material must not cover a large fraction of the membrane surface or cause stagnation in the solutions near the membrane surface. The stack design developed by lonics. Inc.
The ED stack in Figure 8. open channel in the solution compartment. The thick membranes and spacers result in higher electrical resistances than those found in sheet-flow stacks. The Vexa@ non-woven polyolefin netting developed by DuPont has gained wide acceptance as a spacer that meets these criteria. The spacing between membranes of sheet-flow stacks is made as thin as practical to reduce electrical resistance and minimize the thickness of hydraulic boundary layers at the membrane surfaces. SPKEI Fiu. but the openness of the solution channels should make them less prone to fouling by suspended solids. Yet another stack design utilizes unit cells.3:
Main components of an electrodialysis stack. the membranes made by lonics are substantially thicker than those produced by most other manufacturers. Therefore. The envelopes also have
. so a more rigid membrane is required to bridge the space without severe deflection.5 to direct solutions over a relatively long.3 represents the sheet-flow type. The sheet-flow design is used by most ED manufacturers (lonics. The flow of solution in the plane of Vexar netting is illustrated in Figure 8. is quite different from the sheetflow stacks previously described. AND ANOTHER EN0 FRAME
-
_
E -
Figure 8.

The electric current carries ions from the external solutions through the membranes to the inside of the envelopes where they are trapped. For small installations where the throughput does not justify multiple stacks. Recycle is inherently less efficient than once-through flow in both stack utilization and in energy consumptions since the same solution must be pumped and desalted repeatedly and then remixed with a more concentrated solution. The entire assembly of alternating membrane envelopes and spacer screens is held between a set of electrodes. The concentrated solutions inside of the concentrate cells flow through small tubes that lead from inside the concentrate cells to a plenum chamber arranged outside the stack. Thus. In ED. water with 2.000 ppm TDS (total dissolved solids) could be reduced to 500 ppm TDS by two stacks that each achieve 50% desalting. and some type of staging is needed for further desalting.
APPLICATIONS
OF ED
Outside of Japan.000 ppm. the largest application of ED has been in the desalination
. Only osmotically and electro-osmotically transferred water flows through the membranes. Then a second stage with 50 cell pairs would reduce the salinity to 500 ppm. feed and bleed.490
Handbook of Industrial Membrane Technology
spacer screens between them to insure uniform solution flow across the membranes. and ion exchange the necessary degree of desalting can often be achieved with one pass through the device. In large installations such as municipal water supplies this staging is achieved by passing the water through a series of stacks. For the example cited above. However. Ten-fold enrichment or 90% recovery of some feed waters as diluate can be achieved by concentrate recirculation if solubility limits of the dissolved substances are not exceeded in the concentrate stream. The first stage would utilize 100 cell pairs to reduce the salinity to 1. In evaporation. there are other alternatives including internal staging.
STAGING
OF ED STACKS
ED differs from other desalting processes in the degree of desalting achieved in a single stage. reverse osmosis. Feed and bleed is useful for operations where the enriching stream needs to be as concentrated as possible. but with half the TDS reduction because the water throughput per cell would be doubled. the degree of desalting is usually limited to about 50% per pass. and batch recirculation. Each cell pair in the second stage would remove the same quantity of salt as a cell pair in the first stage.000 ppm TDS. the increased flexibility in process control makes recycle systems attractive for small-scale operations where stacks are over-sized to handle varying loads. For example. Internal staging allows one to utilize one pair of electrodes to achieve multi-stage desalting. one might use a stack with 150 cell pairs to treat water with 2. With feed and bleed or batch recycle systems some or all of the water that has already been processed by the stack is mixed with feed and returned via a recirculation pump. unit cells achieve the maximum degree of concentration that is possible with ED.

4 ED desalting of seawater has also been demonstrated on a commercial scale. Its greatest use has been in whey deashing. the liquid streams had to be cooled.A. Ion exchange has traditionally been used for preparing waters with low salinity. The inevitable accumulation of solids on the membranes and spacers required routine cleaning in place and periodic disassembly of the stacks for mechanical cleaning.Electrodialysis
491
of brackish water. and the Office of Saline Water in the U.’ Still the greatest use of ED treatment of seawater has been for the recovery of NaCI. and it contains useful quantities of proteins. These operations are now well established and can be applied to the treatment of other food products. Japan has no natural deposits of NaCl and had formerly obtained the salt by evaporation of seawater. TNO. lactose.6 Food processing provides many potential applications for ED because of its ability to separate electrolytes from non-electrolytes. Moreover. but the cost of regenerants and the magnitude of the waste disposal problem are proportional to the salinity of the feedwater. The approach has been so successful that essentially all the table salt in Japan is prepared from ED enriched brines. ED might appear to be at a competitive disadvantage with RO for this application. Whey is the waste product from cheese making. ED stacks operate trouble-free for years with little maintenance. because RO removes particulates and substantial amounts of organic solutes from the water. The system shown schematically in Figure 8. the Netherlands National Research Organization. Therefore. Most of the early research and development effort had the objective of improving marginal waters to make them potable. that subject will not be stressed in this chapter. To retard spoilage. The bulk of the dissolved solids can be removed more economically by ED or RO. These two processes are competitive in cost. The use of ED to produce a concentrated brine up to 20% TDS has substantially reduced energy costs (less than 200 kwh per ton of salt). the South African Council for Scientific and Industrial Research. Accumulation of these contaminants on membrane surfaces can cause premature failure of RO modules. Since ED treatment of brackish water is well known. funded programs to improve membranes.6 has been in service at Bell Laboratories since 1975 removing salts. the high mineral content makes it unacceptable for human consumption and of marginal value as animal feed.
. The power consumption has been held to tolerable levels (28 watt-hours/gal) by the use of ion-exchange membranes with low resistance. it is common practice to remove organics and particulates before the rough desalting. and lactic acid. principally CaS04. However.S. Deashing by ED upgrades the whey so that subsequent processing can produce edible whey solids. the use of membranes that are selective for univalent ions has improved salt purity to 97% and reduced precipitation problems. equipment and processes to produce potable water more economically. Whey processing required several modifications to ED systems. and both offer the advantage that they do not contribute additional water pollutants. A potential growth area for ED is in the rough desalting of water that will be subjected to subsequent purification for use as boiler feed or rinse water in the electronics industry. During the 1950’s. from water for use in rinsing integrated circuits. With such pretreatment.

1~6”
HP PUMPS IO-20 PPM TOS P
A I P I
I
LPP%S
RAFWATEA __ INLIZ. The opposite effect was observed in blood ultrafiltrate. Such a process could be used to treat certain diseases where excessive amounts of undesirable proteins are present in a patient’s blood or to obtain needed proteins from a donor. With the apparatus shown in Figure 8. As in electrophoresis.7. Chilling aided in the precipitation of proteins which were removed by ultrafiltration. a solution of such a material at its isoelectric point can be desalted by ED with minimal loss of the material. Other separations of proteins and amino acids have been achieved by pH adjustment of the solution being treated. Desalting of cows’ milk by ED allows larger quantities of cows’ milk solids to be used for these purposes. 200400 PPM TOS
an-f-i /
1 I
1 I
CHARCOAL
BEOS
ELECTROOIALYSIS
UNITS
STORAGE
TANKS
1100 GALS EACH
Figure 8.
. However.6: Hill. a molecule migrates as an anion at a pH above its isoelectric point. (Bell Laboratories. and as a cation below its isoelectric point. NJ)
integrated circuit process water system. Therefore.’ The lumpy texture of thawed frozen milk has been attributed to clumping of micellar casein. and calcium removal led to the dissociation of micellar to serum casein.492
Handbook of Industrial Membrane Technology
LOCAL POLISHING CENTER POINT OF USE
SPECIAL FEE0
AESU cANlST. The salt content of the plasma was restored by passing it through the concentrate compartments of the ED stack before returning it to the host.. Jain* used ED to desalt the blood plasma. Murray
Cows’ milk is more salty than milk from human mothers. Research has shown that desalting by ED to remove calcium improved the protein stability of frozen skim milk and its concentrates. a molecule that is at its isoelectric point will not migrate in an electric field.ERs_ -----==z====dl
NUCLEAR GRADE CANISTER
NUCLEAR GRAD_ F RESIN CANISTERS
MiXF” . and this limits its use in the preparation of infant formula.

8.8:
RINSE WASTE
ED for recovery of nickel from electroplating waste water. The used rinse water flows through the depleting compartments of the ED stack where the metal ions are transferred into the concentrate stream. The rinse streams from such processes pose particularly troublesome pollution problems. They are usually too dilute for direct metal recovery and too concentrated for disposal.
-
A I: iii 0
1
I I
COOL
-
1
fro
I ’
HEAT 0 -IUF I I -
Figure 8. ED processing of a rinse stream from a nickel electroplating system9 is illustrated in Figure 8.
ANTICOAGULANT
HOST
UF
. The concentrate stream can be recirculated to build up its metal content to a level that is useful for further recovery or direct return to the plating bath.
.
RECOVERED PLATED a--4 I PARTS i---r t
RNSE TANK
L
Figure 8.7:
Membrane system for recovering proteins from blood. The treated rinse water can then be reused in the process.Electrodialysis
493
Metal finishing processes offer numerous applications for ED in pollution control and material recovery.

lo The process illustrated utilizes an uncharged EDTA-CU complex where the reduction of copper is driven by the oxidation of formaldehyde to formic acid. Nevertheless. Subsequent elevation of the potential in the latter case will increase the current which is carried by the only available ions.9 for the electroless plating of copper onto nonmetallic substrates such as printed circuit boards. NaOH is added to adjust pH sufficiently to ionize the formic acid. and the power consumption is about half that of electrolytic cells. A similar approach was used to recover silver from spent photographic bleach-fixer solutions. bipolar membranes offer the prospect of low cost and minimum unwanted by-products. the current will be relatively high. 1 N) acids or bases are needed. remains in the bleach-fixer solution. because Donnan exclusion diminishes with increasing solution concentrations. Compared to the electrodes used in conventional electrolytic cells. the iron. The current efficiency of acid/base generation and the purity of the acid and base made with bipolar membranes drops off as concentrations increase. A commercially available ED stack with bipolar. The formate ions and the sulfate ions introduced with the make-up copper are undesirable reaction by-products that are removed by ED. the energy associated with gas evolution is saved. but the amount of current flow will be determined by the orientation of the membrane.. Since there are no gasesevolved at the bipolar membranes. which was used to oxidize the silver. which is returned to the plating tank.10. H+ and OH. The silver permeates the cation-exchange membranes to the iron-free concentrate stream from whence it is reclaimed by electrodeposition in a separate cell. the current will quickly drop to a low value as the anions and cations are depleted from the membranes. If the anion-exchange membrane faces the anode. If the cation-exchange membrane faces the anode.11. the production rate is limited by the rate of diffusion of water into the bipolar membrane.
.ions generated at the junction by water splitting.494
Handbook of Industrial Membrane Technology
ED has also been used to improve or maintain the quality of a plating bath and thus eliminate the need for periodic replenishment. Application of an electric potential of about one volt to the bipolar membrane exposed to an electrolyte solution wil! cause current flow. Where dilute (e. This results in the production of acidic and basic solutions at the surfaces of the bipolar membranes. An instructive example is shown in Figure 8. A current density of 100 mA/cm’ was maintained with an applied potential of 2V/cell. The uncharged EDTA-CU complex remains in the diluate stream. there are substantial advantages to the process.g.” As shown in Figure 8. as illustrated in Figure 8. Further. because it forms a complex with EDTA. anion and cation membranes was used by the author to generate 1 N HCI and NaOH from NaCI. Multiple bipolar membranes can be placed between a single pair of electrodes in an ED stack along with other ion-exchange membranes for the production of acid and base from a neutral salt. the bipolar membranes are inexpensive.
BIPOLAR
MEMBRANES
Bipolar membranes consist of an anion-exchange membrane and a cationexchange membrane laminated together.

12 shows a process for absorbing SOs from flue gas with a solution of NasSOs and regenerating the solution with a water splitter containing bipolar membranes.496
Handbook of Industrial Membrane Technology
o
t-
Figure 8.
. The transfer of electrical charge from electronic to electrolytic conduction is accomplished via electrode reactions.13. Current flow in the circuit external to the stack is electronic. current flow within the stack is electrolytic. An ED stack with alternating bipolar and cation-exchange membranes (the arrangement shown in Figure 8. This potential is applied from an external power supply through electrodes situated on either end of the stack of membranes and spacers.e. i. H+ ions generated in the bipolar membrane lower the pH of the solution and allow SO2 to be stripped out.13) was used to convert Na2C03 and/or NaHCOs to more valuable NaOH with gaseous CO2 as the by-product.
The current efficiency was about BO%.e. Several applications for bipolar membranes have been reported.13 A stack with bipolar.igure 8.. and the purity of the acid and base exceeded 98%. Figure 8. more applications of bipolar membranes are sure to follow. Na’ ions pass through the cation-exchange membrane into the basic solution. i. The stripped acidic solution is combined with the basic solution and returned to the absorber.11:
Bipolar membrane construction and operation.12 The regeneration process is shown in F.14 As more bipolar membranes become available in commercial quantities..
ELECTRODES
AND STACK POWER
Electrode Reactions and Materials An electric potential is the driving force for transport of ions in ED. anion-exchange. and cation-exchange membranes was used to convert ethylenediamine dihydrochloride into ethylenediamine and HCI in separate streams. electrons flowing through metallic conductors. However. ions flowing through solutions.

While acid addition contributes significantly to operating costs of commercial ED plants. The presence of pH-sensitive salts. Indeed. lead and iron. fouling problems can be severe at the cathode.498
Handbook of Industrial Membrane Technology
Anode Reactions
MO --f M+X + xe-
metal dissolution (acidic solution)’ (basic solution) evolution of > gaseous oxygen
2Hz0 40H--+
-+ 0s + 4H+ + 4eO2 + 2Hz0 + 4e-
2CI. titanium is not suitable for reversible electrodes..g. Moreover.ions generated at the electrodes can cause undesired reaction with the components of the feedwater and lead to membrane fouling. platinized niobium has been used. is used to calculate the amount of acid required to neutralize the OH. because the OH. Electrode Isolation The reactions that occur at the electrodes are normally incidental and occasionally detrimental to the desired separations that take place in the repeating cell pairs of an ED stack. In contrast. The use of a thin layer of platinum on a substrate of less expensive metal (e. Ca(HC0-J2 in the feedwater is a matter of concern.g.500 A-sec/eq.. because it is attacked by the Hz generated at the cathode. The H+ and OH. Although the reactions at the cathode are not severe in terms of electrode deterioration. Faraday’s constant.” When the polarity of the electrodes must be reversed periodically.+ CaC03 + COT + 2HsO The problem of pH control in the cathode rinse stream (catholyte) is usually handled by acidification. reversibility of ED electrodes is exceptional. titanium) moderates the expenditures for anodes.tially. Some metals. However. Care must be exercised to ensure that the polarity of ED electrodes is not inadvertently reversed. migration of some feed components into the electrode rinse streams can cause precipitation on the electrodes or the spacers and membranes close by. because severe electrode deterioration can result. acidification of the feed stream to prevent precipitation can increase operating costs substar. e. but they are not reversible. 96.. will form stable conductive oxide coatings under properly controlled anodic conditions.+ Cls + 2eMO +xOH2M” + 2xOH-
evolution of gaseous chlorine oxidation of electrode
-+ M(OH)x +xe+
MsO. Therefore. Ca(HCOs)s + 20H.
. e.ions generated at the cathode. as in the EDR process.g. the fact that there are hundreds of cell pairs between each pair of electrodes keeps the catholyte acid costs at tolerable levels.ions generated at the cathode can lead to precipitation. it is prudent to carefully consider such possibilities and isolate the electrodes to minimize undesired reactions. + xHzO + 2xe-
Anode reactions that produce dissolution or oxidation of the electrode metal are avoided by selection of resistant metals such as platinum.‘6”8 Metal electrodes with oxide coatings are economical compared to those with platinum coatings.

With a 2. so neutralization of the anolyte may be needed if acid-sensitive components.g. but they are critical issues in large ED stacks.20 A similar arrangement with NaOH solution for the electrode rinse allowed the use of an inexpensive ferrous anode in an ED stack. The simplicity and economy of a single electrode stream must be weighed against the problems of current leakage through the solution lines. The use of a univalent-cation-selective membrane next to the cathode can alleviate some of the problems of multivalent cations in the catholyte. .5% suspension of carbon in 0. The use of cation-exchan $ membranes at both electrodes made the anolyte a depleting stream and the catholyte an enriching stream. a current of lo-20 mA/cm’ could be sustained for weeks with no gas evolution. The undesirable formation of hypochlorite occurs at the anode whenever chloride ions are present in the anolyte. The overriding advantage of the cation-exchange membrane is that it blocks the entry of cathode-generated OH.ions into adjacent compartments. An anion exchange membrane is often used to block the passage of H’ ions into the rest of the stack. anion exchange membranes are only moderately effective at blocking H+ ions. which presents disposal problems. Na2S04) for the anolyte. The major advantage to this arrangement is that cations from the feed solution can be almost completely excluded from the catholyte.g. e.14. and the mixing of H. proteins.Electrodialysis
499
Selection of the type of membrane to be used at the cathode end of the stack is important.out of the electrode rinse solution and conserved the Na&Os that maintained the conductivity in the electrode compartments. Kedem et al demonstrated that gas evolution could be avoided by the use of powdered activated carbon in a common electrode rinse stream. but this arrangement makes the anolyte a depleting stream and requires a supply of electolyte (e. makeup water of good quality is required. These problems may be conveniently ignored in laboratory experiments. the catholyte is an enriching stream. This charge was neutralized by sorption of anions from the rinse solution. However.02 to 0. The catholyte is a depleting stream if an anion-exchange membrane is placed next to the cathode.. the charge on the particle was reversed and the sorbed anions were exchanged for cations. buildup of undesirable materials such as hypochlorite.2 normal solution. NaOH.lg The large surface of a carbon particle became charged when it came in contact with the anode or with another positively charged carbon particle. If a cation-exchange membrane is used. The benefits of .. are present in the feed. and O2 generated at the electrodes.a common electrode rinse stream were demonstrated in a photographic devel per regeneration process illustrated in Figure 8. The use of a cation-exchange membrane at the anode prevents chloride entry from the feed solution to the anolyte. When the suspension reached the cathode. In some circumstances.. The price to be paid is that acid must be added to the catholyte in sufficient quantity to supply all of the anions needed to carry current through the membrane. a common rinse solution can be circulated through the anode and cathode compartments. An enriching catholyte stream is almost invariably a waste stream. Moreover. This kept the 8r. The major considerations about isolation of the anode pertain to the acidic. oxidizing conditions there.”
. This means that cations from the feed solution can enter the catholyte.

15. The output from the isolation transformer has a floating ground if the load is not intentionally grounded. An isolation transformer contributes measurably to the safety of operating an ED stack. especially an experimental one from which samples are being taken. A full-wave rectifier is commonly used to convert AC to DC. voltage adjustment must occur before conversion to DC. Since transformers operate on AC.500
Handbook of Industrial Membrane Technology
DEVELOPER
Figure 8. The components required to safely deliver current (DC) at the proper voltage are illustrated in Figure 8. In laboratory ED units an autotransformer or a solid state voltage regulator is preferred.14: The use of cation-exchange membranes at both electrodes conserved the Naz CO3 in the common electrode rinse stream of a photographic developer regeneration process. If the circuit is accidentally grounded by an operator’s hand. The power level of the ED stack is regulated by adjustments of the stack voltage. The rates of build-up and collapse of concentration boundary layers are so slow compared to the ripple frequency that stack performance is unaffected by the ripple.
. The current from such a rectifier has a 120 Hz ripple which is of no consequence. or the number of cell pairs in the stack may be varied to match the desired cell-pair voltage drop to the available DC voltage. because the user will likely want to experiment with the effects of applied voltage on stack performance. In large installations a transformer may be used. this contact point becomes the zero potential point.
Power Supplies Ionic transport through the solutions and membrane of an ED stack is driven by a direct electric potential. but no shock is experienced.

Otherwise. With constant R. E.
Connections between the power supply and the electrodes are usually made outside the stack so that they can be kept clean and dry. because I is proportional to concentration..Electrodialysis
A-C SOURCE
501
ISOLATION TRANSFORMER
LJJ
AUTOTRANSFORMER
FULL-WAVE RECTIFIER
METERS
DRY CONNECTIONS TO ELECTRODES
Figure 8. varies inversely with concentration. I a C. and R.. severe corrosion problems can result. It is advisable that the conductor between this connection and the electrode be of the same material as the electrode and that the connection with the electrode be welded.16) over the range of safe operation of ED allows one to develop a simplified equation concerning power consumption in an ED stack. Since power consumption is the product of current and voltage. i. E.e. the current is proportional to the amount of salt removed by the passage of current. i. = IR.. does not change greatly as feed concentrations change. The observation that the cell-pair resistance is reasonably constant (see Figure 8.. The typical voltage drop across a cell pair is in the range of 1 to 2 volts. the cell voltage is proportional to cell current.e.. Assuming constant current efficiency...15:
Power Supply for ED. one reaches the not-too-surpris-
..

1.94-0. a simplified mathematical model will be used to show the quantitative aspects of concentrations polarization and explain the important concept of “limiting current density”. I
I I
6.0 z
-
0 * -1
I.502
150. it was shown that the interposition of pairs of anion. * I 900 OR AMP 1 1000 CMIEQ 1200 1400
0
1
16 00
Figure 8.97
EFFICIENCY
110 -
100 -
B-O-
I I
I
7. the concentration is higher at the membrane surfaces
. P a AC.18: effluent.. Concentration polarization will be discussed first in a descriptive manner.P) and pH of the depleted
ing conclusion that the power consumption in the ED stack is proportional to the amount of salt removed.e. i.0 C. Close examination of the system would reveal nonuniformity of electrolyte concentration within a solution compartment.
Effects of i/N
on cell-pair resistance (R. z
130 -
120 -
COULOMB 0.. but a consequent reduction in electrolyte removal would occur. Then. because it ultimately limits the rate at which ion transport can occur.0167 19. In the enriching compartments.
Handbook of Industrial Membrane Technology
140 -
FF IV STACK ASSEMBLED WlTH TOKUYAHA “EHBRANES AND 23-MIL VEXAR SPACERS VELOCITY 0.and cation-exchange membranes between electrodes resulted in the depletion or enrichment of electrolyte in the alternating solution compartments when an electric potential was applied. 200 400 600 i/E.
.6 E CM/SEC ~aC1 LIMITING FROU VALUE OF COWAN-BRDUN i/y TAKEN PLOT FEED -
-= ” g . Reductions in power consumption can be achieved for any ED stack by reducing the applied voltage. In Figure 8.
CONCENTRATION
POLARIZATION
The phenomenon of concentration polarization is important in electrodialysis.

The flow of electrical charge in this region is designated by the symbol i with the units mA/cm’. Transport numbers within the membranes are not as easily determined as those in solutions because of interactions between the mobile ions and the fixed ionic groups. Solutions with more than one salt or with multivalent ions require a more generalized equation to describe transport numbers: = xizi2ci/l: 1 izi2ci
ti
where Ci is the molar concentration and Zi is the ionic charge of each species present in the solution. The resistance to the flow of electric current increases as this interfacial concentration decreases.9 when external solutions are dilute (<l molar). Now.Electrodialysis
503
than in the bulk solution. consider the region of the system designated by the square in Figure 8. Since most of the ions available to carry current within the membrane are counter ions. and in the depleting compartments. Although the generalized equation above is applicable to the membrane as well as to the solution phase.17 for a single uni-univalent electrolyte such as NaCI. These idealized boundary-layer conditions are obviously unrealistic. the concentration is lower at the membrane surfaces. the concentration of counter ions (those of opposite charge to the fixed ions) is always much higher than the concentration of co-ions (several orders of magnitude higher when external solutions are dilute). The fraction of the total current carried by a particular ionic species in an electrolyte solution is called the transport number t. flux of cations through the cation exchange membrane is proportional to the electric current. With the additional assumption that diffusion is negligible within the membrane. Je+ = i t+/F
Here the overbar designates the membrane phase. This model. the values of h and C are different in the membrane. The F is Faraday’s constant. but they allow a simplified approach to an otherwise complex problem. typically above 0. the transport numbers of counter ions are high. Values of h extrapolated to infinite dilution are shown in Table 8. is based on the simplifying assumptions of stagnant boundary layers of constant thickness and a well-mixed region in the center of the solution compartment.17. What happens with increasing ED current can be presented mathematically with the aid of a simplified model called the Nernst idealization. which is illustrated in Figure 8. Because of Donnan exclusion.2. The reason for the nonuniformity of concentrations in the solution compartments is that cations carry virtually all of the electric current in the cation exchange membranes but only about half of the current in the solution. For a solution containing a univalent electrolyte like NaCl or KNOa the transport number of the cations is:
t
+
= A
+/( x++x-)
where h is the equivalent ionic conductance. especially in the presence of the spacer screens used in an ED stack.
. and i’ is the transport number for cations in cation exchange membrane.

Static N. one can write the steady-state flux equation for cations in the vicinity of the membrane solution interface. Now in the aqueous phase near the membrane.17: Concentration gradients in electrodialysis.4 of an equivalent of cations for each equivalent of cations transported through the cation exchange membrane. the balance would be 0. one equivalent of cations is removed through the membrane for every faraday of current flowing through the region of interest..6 faraday of current would be transported by anions that are moving in the opposite direction. Now. the electric current would deliver only 0. the flux of cations transported electrically is: Je+ = i t+/F If the electrolyte were NaCl where-t’ = 0._
505
-i-i-
-i-T
6 = Static boundary-layers
C = Cation-exchangemembrane A = Anion-exchangemembrane Figure 8. $+ i?/F = = it+/F Jet + Jd + D(Cb-Ci)/ 6
.-ci 1/ 6
The subscripts b and i refer to the bulk and interfacial concentrations. Since i’ = 1 in the ideal case. The net result would be a depletion of anions and cations from the solution at the membrane surface. This depletion would result in a concentration gradient that would allow diffusive transport of the balance of ions needed for electroneutrality. The diffusive flux of ion pairs down this linear concentration gradient is: Jd = olc.4.Electrodialysis .500 A-sec/eq.
96.6 of an equivalent per faraday of current flow. The remaining 0. In this case.

Those data were obtained from electrodialysis of a solution of NaCl in a stack containing ten cell pairs.l However. the value of R. indicating il. and pH of the depleted product water were monitored.r. Moreover. at the cation exchange membrane. R. . the pH of the depleted effluent dropped precipitously. The maximum point in the pH curve was attributed to attainment of the limiting value of i/N at the anion exchange membrane.16. the slight increase in pH of the depleting solution suggests that water splitting was not significant. current. The cell-pair resistance. That limit is approached as Ci approaches zero.t)/ln(Ci.. i is an independent variable that can be adjusted by changing the voltage applied to the electrode. 6.t). However./l.. = (V.5 times higher than that at the cation exchange membrane. .e. i. N = (Ci. The equations above indicate that the value of ilk. A more reasonable explanation is that localized depletion occurred at the inflection point and expanded as current increased. If the concentration were to reach zero over the entire surface at the same value of i/N. However. membrane and salt type)... [The polarization parameter is defined as the current density.16 show the results of an experiment to determine the values of irim. A control experiment with one cell pair in the stack was used to determine the voltage drop attributable to the electrodes and rinse streams. the inflection would likely be more abrupt and the curve would become steeper. whether the concentration approaches zero in the solution at the interface of the cation exchange membrane.506
Handbook of industrial Membrane Technology
Rearrangement yields: i/(Cb-Ci
1
=
DF/6 (?+-t+)
F is a constant and D. In fact. The limiting current density can thus be expressed: i I im’Cb = DF/ 6 (t+-t+)
For a solution like KCI where the ionic conductances of both ions are the same. divided by the log-mean concentration of the depleting stream. this is not the case for NaCl where t+ = 0. Again this was probably localized depletion that expanded as current increased.4 and t-= 0. i/N. The ratio of the irim values was 1. because splitting water to form H+ ions would be the only source of ions to carry the additional current. Now.. at the anion exchange membrane would be 1../Co.400 A-cm/eq. When the value of i/N reached about 1.r . was observed to increase slightly for moderate increases in the polarization parameter. one might question whether the inflection points in the curves truly represent the limiting values of the polarization parameter. indicating ir.. flow rate. The data in Figure 8. but the pH decrease was much more dramatic than the pH increase observed when polarization occurred
..and irim is the same value for both interfaces. When i changes.6.Cd.n. t+ = t.V. increased more rapidly when the value of i/N exceeded 950 A-cm/eq.. i = I/A... Upon close examination of the data in Figure 8. a slight increase in the pH of the depleted effluent was detected. there is a limitation on the extent to which the value of i in this equation can be increased.. t’ and t* are essentially nonvariant for a given set of operating conditions (temperature.21 The applied voltage.. at the anion exchange membrane..) x A.5 as expected. the value of Cb-Ci changes proportionately.. i/N.

Such problems are usually handled by reducing current densities. Colloidal particles. these charged molecules tend to accumulate on the dilute side of anion exchange membranes and result in a buildup of membrane resistance.18 illustrates some of the fouling problems.g. Large organic molecules with ionizable groups (e. Fouling of membranes by organic anions and colloidal material
. and by addition of sodium hexametaphosphate to delay crystal formation. Moreover. Therefore. because their concentrations are highest there due to concentration polarization. for a margin of safety. The reaction of HCOs. Consequently. CaS04 precipitation is countered by limiting the concentration in the enriching stream below the saturation level. Also. the presence of non-conductive spacers in the solution compartments causes local current densities to be higher than the calculated average value. This observation indicated that water splitting occurred at the anion exchange membrane but not at the cation exchange membrane. ED stacks can be operated for years with little concern about membrane fouling. one would operate an electrodialysis stack at a current density substantially below the measured limiting value. most feedwaters have constituents that can cause problems for ED stacks. However. Their negative charges allow them to migrate in the electric field but their large size prevents their passage through the membrane.22 Instead substantial water splitting occurs in the anion exchange membrane. CaCOs precipitation occurs when the limiting current density is exceeded. Figure 8.. Uneven flow distribution within or between solution compartments could lead to localized thick boundary layers that result in lower values of ijim. periodic current reversal (the EDR process). CaCOs precipitation can be prevented by acidification of the enriching-steam feed.ions with OH. which are also usually negatively charged. Sparingly soluble salts such as CaS04 and CaC03 can precipitate on the concentrate side of the membrane. This terminology obviously does not refer to a physical limitation to the current flow but rather to a practical limitation for trouble-free operation of electrodialysis. begins to rise rapidly is generally referred to as the limiting current density.Electrodialysis
507
at the cation exchange membrane.16 would indicate. Problems of membrane fouling are generally less severe when current densities are kept low to minimize concentration polarization. The current density at which R. humic acids from decomposing vegetation) are troublesome foulants for ED membranes. cause similar problems. and it is catalyzed by tertiary amines that withdraw protons from water molecules..
MEMBRANE
FOULING
When the feed solutions to ED stacks are clean and are relatively free of sparingly soluble materials. The prevailing theory explains that the kinetics of water splitting in solution are too slow to cause appreciable pH changes to occur. or pretreatment of the feedwater. the precise value of irim is not as easily established as Figure 8.ions generated by water splitting forms CO. ions that pass through the anion exchange membrane and precipitate in the boundary where the concentration of Caf+ ions is the highest.

freshly formed precipitates tend to dissolve. these membrane-fouling problems can be alleviated to some extent by use of the EDR process. the extent of pretreatment in EDR plants would likely make the water acceptable for conventional ED. reduced chemical costs. it has not been adopted by other suppliers of ED equipment. but rather to a general impression that removing foulants by pretreatment is more cost effective than attempts to moderate their damage to the membranes. Although polarity reversal has been lauded as a major breakthrough by lonics. since current reversal causes concentrating boundary layers to become depleting boundary layers. This loss of feedwater along with the high costs associated with the automatic switching valves are the major drawbacks to EDR. lnc. Immediately after the current reversal the product water is of poor quality due to the discharge of the aforementioned foulants. higher operating currents. and less maintenance than with conventional ED. there is a period of a few minutes when the product water must be diverted to waste. Supersaturation of CaS04 up to 220% in the brine stream was tolerated by EDR..
Large organic ions and pH-sensitive salts can foul anion-exchange
is best handled by pretreatment to eliminate the offending material. Therefore.” In fact.18: membranes. Alternattively. The supersaturation was pushed to 440% with the addition of acid and sodium hexametaphosphate.24
. Moreover. The reversal is programmed to take place about three times per hour. EDR is the electrodialysis reversal process developed by lonics. This frequent reversal tends to dislodge the organic and particulate foulants from the membrane surfaces. Inc.508
Handbook of industrial Membrane Technology
-
+I
+ Ca++
HCO.
+ + + + + + +I d
Figure 8. This avoidance of polarity reversal does not appear to be due to strong EDR patents. The proponents of EDR23 claim that these disadvantages are more than offset by increased membrane life.23 The EDR process employs reversible electrodes and automatic valves to swap the flows of enriching and depleting product lines.

and pretreatment equipment. Electrochem. S.. Water transport in ion exchange membranes.J.S. Komori. and Watanabe. These costs can be amortized over the expected life of the system and expressed in terms of product output. Desalination 32: 383-9 (1980)... Patent 4.. power supplies.H. 3: 57-70 (1978). Patent 4. New York: McGraw-Hill (1962).. European Patent 0015737 (1980). The costs of desalination by ED and RO are very close..J.
13. 9: 731-6 (1979).
15. and Amundson. D.238. Chang. (1980).. tanks. J. J. Grenda. Proc.. and Nagasubramanian. 5.5 power costs are usually in the range of $0. U. and variable operating categories. .A. Chlanda. 12. Energy-saving electrodialyzer for seawater desalination. I. Except for seawater where consumption of 28 wh/gal was reported. T. Capital costs are usually above $l/gpd capacity except for very large plants.. The costs are generally divided into capital.. because the current is proportional to the amount of salinity reduction.W. Ono.B.R. Y. C.
11. Int.. Tech. building.. F.. Sci. Symp. Desalination 34: 77-95 (1980). T.25 There are economies of scale-up in all of the costs except energy and chemicals.204. 0. 4. Patent 4. Capital cost include the site. ED is generally the less expensive proces? where the salinity is low. J. piping. Proc. These cost begins at about $l/kgal for large desalting plants for desalting brackish water to drinking water standards. Y. 9th.. J.. V. P. and Jenczewski. /on Exchange. ED: Polarity reversal or not? Paper presented at 46th International Water Conference. Foster. Electrical power consump tion varies with the salinity of the feedwater and the applied voltage. Zmolek.J. Trivijitkasem. P.. M.. Rowe. K. M. Assoc. 6. Jain.. Itoi. U..P.. Water Supply Improv. S. 47: 1429-34 (1982). Membr. but they can reach several dollars per kgal for small installations. Gancy. but low-resistance membranes and thinner solution compartments could make ED competitive for seawater desalting.
IO.S. and Farrar. Acta 25: 271 (1980). and Nagasato.50 to $l. 2. $/kgal. T. fixed operating. RO is considered to be less expensive for seawater.. A. Conf. Dee: 6-11 (1977).
REFERENCES 1.140 (1981).Annu. K.275. Paper Number 4 (1982). D. Seto.
14. H. Kneifel.Electrodialysis COST OF ED
509
The costs of constructing and operating an ED system are dictated by the quality of the water to be treated and the degree of desalting that must be achieved. pumps. F. Pennsylvania (November 1985). Liu. Salt. K. Appl. Ind. C.g. K. Pittsburgh. e.
Lonergan.305 (1980). U.
.930 (1980). Water Eng.
8. Fennema. 5th:
317-24 7. Nakamura..M. Electrochim. and Ostvold. Helfferich.S. 9. and Kawahara.B.OO/kgal. 3. T. R. stacks. Tani.. Food&i. Kawate.A. Trade Fair Nat’l. T. and Hattenbach.

The reformed complexing agent then diffuses back across the membrane.1) .e. which is maintained at a lower pH. The water-immiscible agent fills the pores of a microporous membrane.1 for Cu2+ and H+ ions and a complexing agent denoted by RH. forming a neutral-ion complex. the left interface in Figure 9. The equilibrium that exists at the two membrane-solution interfaces is: 2RH + Cu2+ _ CuR2 + 2H+
At the feed solution interface (i. The desired ion is complexed at one interface of the membrane. ano it diffuses across the membrane to the product solution interface.9
Coupled Transport Membranes
Richard Baker and lngo Blume
INTRODUCTION Coupled transport is a membrane process for concentrating ions and separating ions from aqueous solutions. 511
. The coupled transport process is illustrated in Figure 9. thus forming a liquid organic membrane. There. the complexing agent acts as a shuttle to carry ions across the membrane. where it picks up more of the desired ion.. This step reforms the neutral complexing agent. The neutral copper complex CuRz is soluble only in the organic phase. where the reaction is reversed by making appropriate changes in the external solution conditions. and the movement of hydrogen ions in the opposite direction maintains electrical neutrality. the reaction is reversed. Thus. liberating two hydrogen ions. Thus. liberating the copper ion and consuming two hydrogen ions. The membrane used in this process consists of a water-insoluble liquid containing an ion-complexing agent that is specific for the ion of interest. The neutral ion-complex then diffuses across the membrane to the opposite interface. the copper ion reacts with agent. It is this coupling between the flow of one species (CL?‘) and the other (H+) that gives coupled transport its name. copper ions move from left to right. which then diffuses back across the membrane.

’ Perhaps the first coupled transport experiment was performed by Osterhout.512
Handbook of Industrial Membrane Technology Carrier Membrane
Cu*++ SHR-CUR
2
+2H+f
z
CuR
2
+2H+-Cu*++2HR
Figure 9. The energy for the pumping action derives from the flow of one species (hydrogen ion in the example shown here) down its concentration gradient. and workers began to develop synthetic biomembranes analogues of the natural systems. who studied the transport of ammonia across algae cell walls. for example. in the mid-1960’s.can be transported by a completely analogous processes using.
HISTORY
AND BACKGROUND
The historical development of coupled transport is shown schematically in Figure 9. At that time. Some workers are continuing to apply these membranes to metal separations.1:
Coupled transport scheme for copper. ions of interest can be concentrated against their concentration gradients. Pfeffer postulated transport properties in membranes using carriers. shown in Figure 9. the PVC/ester film was cast on a paper support. Coupled transport is not limited to cations.‘* but most current interest in PVC matrix membranes is in their use in ion selective membrane electrodes.“. tertiary amine complexing agents. Researchers actively pursued this work until the late 1960’s. Bloch and Vofsi prepared immobilized liquid films by dissolving the esters in a PVC matrix. Further.3.2. As early as 1890.‘3
. Typically. For example. interest in this approach lagged. Bloch and Vofsi published the first of several papers in which coupled transport was applied to hydrometallurgical separations. apparently because the fluxes obtained did not make the process competitive with conventional separation processes. the carrier concept was well developed.3 By the 1950’s. Coupled transport membranes can thus be considered as chemical pumps. Sollner and Shean4” studied a number of coupled transport systems using inverted U-tubes. At the same time. Anions such as Cr207*.’ This process originated in early experiments of biologists using natural carriers contained in cell walls.
Proper selection of complexing agent and conditions makes clean separations of metal ions possible. namely the separation of uranium using phosphate esters7-” Because phosphate esters were also plasticizers for polyvinyl chloride (PVC).

emulsion droplets are separated from the feed and the emulsion is broken. The organic carrier phase forms the wall of the emulsion droplet separating the aqueous feed from the aqueous product solutions. the emulsion membranes must be completely stable during the extraction step.
Feed
Solution
Product
Solution
Microporous Polymeric Membrane
Figure 9. When sufficient metal has been extracted. Metal ions are concentrated in the interior of the droplets. to prevent the two aqueous phases mixing. In this technique.
The second type of immobilized liquid carrier is the emulsion or “bubble” membrane. the first pilot plant was not installed until 1980. but must be completely broken and easily separated in the stripping step. In the first approach. the liquid carrier phase is held by capillarity within the pores of a microporous substrate. Ideally.4. as shown in Figure 9.” The principal objective of this early work was the recovery of copper and other metals from hydrometallurgical solutions. Despite considerable effort on the laboratory scale. The technique of emulsion membranes was popularized and fully developed by Li and his coworkers at Exxon. This approach was first used by Miyauchi14 and further developed by Baker et al15”7 and by Largman and Sifniades. liberating a concentrated product solution and an organic carrier phase. the Exxon group’s work led to the installation of the first pilot plant in 1979. Achieving this level of control over emulsion stability has proved difficult.5. a surfactant-stabilized emulsion is produced as shown in Figure 9.‘9”o The principal problem is instability of the liquid carrier phase in the microporous membrane support.2’-2g Starting in the late 1950’s and continuing for more than twenty years. a number of pilot plants have been installed.29 Although the process is still not commercial.4:
Organic Carrier Phase
Supported liquid membrane.514
Handbook of industrial Membrane Technology
Following the work of Bloch and Vofsi. principally on hydrometallurgical
. two other methods of producing immobilized liquid films were introduced. The principal technical problem is the stability of the liquid membrane. The carrier phase is decanted from the product solution and recycled to make more emulsion droplets. Both are still under development.

depending on the type of reaction occurring between complexing agent and permeant. and Danesi et al.41r42 Martin and Davis43 in the United Kingdom.6). Marr and Kapp in Austria.30-34 More recent workers in the field include Halwachs and Schuger3sJ8 in West Germany. aa.
...(. Another important group working independently on the problem at about the same time was Cussler. . _ . Christensen and lzatt48”’ in the United States..s2”s
Organic Carrier Phase
Product
Solution
_ _(.and A is Sod*-. The first type is called counter transport (shown in Figure 9. Evans and others at Carnegie Mellon.5: Surfactant stabilized emulsion membrane. ” \. For example. (R3NH)2S04 + Hg2(Sg4)2 --(R3NH)2U02(Sg4)2 In this case.. consider the transport of uranyl ions by the tertiary amine salt (R3NH)*S04. the reaction can be written nRA + M _ MRn + nA (3) CuR2 + 2H+ (2)
In the reaction for the transport of copper..<.
COUNTER
TRANSPORT
AND CO-TRANSPORT
Coupled transport processes can be divided into two categories.39t40 Stelmaszek et al in Poland.. + Sg42(4)
. I .Coupled Transport Membranes
515
feed streams. M is UO2(SO4)*.46f47 and Lamb. The key feature of counter transport is that the fluxes of the two permeating ions move counter to each other across the membrane. M is Cu*+ and A is H’. _ .. Equation 3 does not have to be limited to cation transport. lrJl \ Feed Solution
Figure 9. s. The reaction in this case is: 2RH + Cu2+ Z=? More generally..44t4s Noble et al. A number of reviews on various aspects of liquid membranes and coupled transport processes have also appeared.

7. illustrated in Figure 9.
The second type of coupled transport is co-transport.6:
Counter-coupled transport. The general form of the reaction in this case is: A+N+R -RAM (5)
A typical example of this type of process might be the transport of uranyl ions by tertiary amine complexing agents via the reaction: 4R3N + 4H+ + 002(SO4)3 ‘1’=(R3NH)4 002(SO4)3 (6)
. The key feature of co-transport is that the fluxes of the two permeating ions move in the same direction across the membrane.516
Handbook of Industrial Membrane Technology
Membrane
nRA + M--MRn
+ nA
MRn + nA-+nRA
+ M
---w
MR /
M A
I
Dilute A and M M --mA
I t
/
/AR \
Concentrated A and M
4
M A
Figure 9.

THEORY Several authors have attempted to provide a general theory of coupled transport and the closely related process of facilitated transport.”
Membrane
A + M + R-rRAM
RAM-A
+ M + R
I
I
M A
Figure 9.Coupled Transport Membranes
517
In this example. Another example of co-transport is the transport of alkali metals by crown ethers by reactions of the type: Cl+ Na+ + Crown -Crown (NaCl) (7)
where Crown is a cyclic ether. M is UOz(SO&4.7:
I t t
l
Co-coupled transport.and A is 4H’. including Smith
.

all four species are in equilibrium in the aqueous solution and in the organic membrane phase. but in practice CuRz and RH are confined to the organic membrane phase and Cu*+ and 2H’ are confined to the aqueous phase.44r45 The usual approach is to assume an equilibrium between the various reacting species in the membrane phase and link these concentrations by the appropriate equilibrium constants. .
(8)
This reaction is characterized by an equilibrium constant:
K =
CMRnI [Al”
[RA]” [M]
(9)
This equilibrium constant could be written for either the organic phase or the aqueous phase. superscript ” refers to the aqueous phase. K. We therefore assume that concentrations of the permeants at each interface to be defined by an appropriate equilibrium constant and then substitute these values in a simple Fick’s law expression. write Equation 9 as: [ MRn] ’ [A] ‘In K’ = [RA]‘” [Ml” = km k. where [A] and [Ml are negligibly small. These expressions are complex and of little practical value. consider the counter transport reaction: 2RH + Cu2+4 *CuR2 + 2H+
.] and [RA] are negligibly small. For example. Equilibrium between the reactants exists only at the membrane interface. are the partition coefficients of M and A between the aqueous and organic phases. The same equilibrium constant must apply at both membrane-solution interfaces. only [A] and [Ml are measurable in the aqueous phase.518
Handbook of Industrial Membrane Technology
et al. and k.
(IO)
where the superscript ’ refers to the organic phase. inside the liquid membrane: nRA + fl _ MR n+nA
. therefore. The result is an expression for the permeant flux that includes terms for permeants diffusion coefficients. and k. For example. A simpler approach others have followed15’16!60 is to ignore the effect of minor species in the membrane and aqueous phases. only [MR. We prefer this formulation of Equation 10 because all the quantities are easily accessible experimentally.56 Schultz et al5749 Cussler5’ and Danesi. and we can recast Equation IO into the following form:
.] and [RA] are measurable in the organic phase. and distribution coefficients. where [MR. [M&l '/D'Jl ” is easily recognizable as the distribution coefficient of metal between the organic and aqueous phases.
In principal. equilibrium constants. Similarly. Consider the general case of a permeant M of valence n reacting with carrier RA to form the complex MR. We can. However. A diffusion model using Fick’s law experiments is then set up for each species and these expressions are integrated between the boundary values defined by the membrane distribution coefficients and the permeant concentrations in the surrounding aqueous solutions.

(11)
where o and II refer to the two sides of the membrane. At steady state. To illustrate coupling effects.] b and [RAI b” = [ RA] h”.~~]*~~l+l)o ([A]l~/[:I”K1+1) J (15)
This shows the coupling effect. However. in mol/cm2-see. and we obtain the expression:
(12)
Thus. JMR .!..Coupled Transport Membranes
519
K’
=
CMRnIi [All”
[RA]$-’ [Fl]b: =
CMRnIi CA]&”
[RA-jkn [M]. In principal. the maximum concentration factor of metal ion that can be established across the membrane varies with the counter ion concentration ratio (in the same direction) raised to the nth power. Thus. However. Under this condition. demonstrates nothing about the metal ion flux across the membrane under non-equilibrium situations. can be measured experimentally in a permeation experiment.. across the liquid membrane is given by: 6’lRn ([MRn]. The solution is simple only when n = 1. Consider now the situation where a counter ion concentration gradient is established that exactly balances the metal ion concentration gradient. where DMRn = II is the mean diffusion